Chapter 8

Reactor-Separator-Recycle Networks

8.0 OBJECTIVES

The presence of at least one chemical reactor and one or more separation sections for theseparation of the effluent mixture leaving the reactor(s) characterizes many chemical processes.In almost all cases, one or more of the streams leaving the separation section(s) is (are) recycledto the reactor. In Chapter 6, the design of reactors and reactor networks was considered withoutregard for the separation section(s) and possible recycle there from. Chapter 7 was concernedwith the design of separation sections in the absence of any consideration of the reactor section.Chapter 5, which dealt with the synthesis of the entire process, included a few examples of theinteraction between the reactor and separation sections. This chapter extends that introduction togive a more detailed treatment of reactor-separator-recycle networks.

After studying this chapter, the reader should

1. Be able to determine the best location for the separation section, either before or after the reactor.

2. Understand the tradeoffs between purge-to-recycle ratio, recycle ratio, and raw material loss, when dealing with inert or byproduct chemicals that are difficult to separate from the reactants.

3. Understand the need to determine the optimal reactor conversion, involving the tradeoff between the cost of the reactor section and the cost of the separation section(s) in the presence of recycle, even when chemical equilibrium greatly favors the products of the reaction.

CD-8-14. Understand the conditions under which the recycle of byproducts to extinction can be employed to reduce waste and increase yield.

5. Be aware of the snowball effect in a reactor-separator-recycle network and the importance of

designing an adequate control system, which is presented in Sections 20.3 (Example 20.11) and 21.5 (Case Study 21.3).

8.1 INTRODUCTION

The feed to a reactor section of a chemical process almost always is a combined feed consistingof a fresh feed mixed with one or more recycle streams, as shown in Figure 7.1. Fresh reactorfeeds rarely contain only the reactants for the desired reaction. Besides the reactants, they maycontain inert chemicals, potential reactants for side reactions, catalyst poisons, and products ofthe desired reaction(s). Recycle streams are intended to contain only unconverted reactants of thedesired reaction(s). However, more commonly, recycle streams also contain products of thedesired reaction(s), products of undesired side reactions, and inert chemicals.

Reactor effluents are almost never products that meet purity specifications. Besides theproducts, effluents may contain reactants, inerts, products of undesired side reactions, and feedimpurities. Thus, almost every chemical process that involves a chemical reaction section alsoinvolves one or more separation sections in addition to one or more recycle streams. A majorchallenge of process design is to devise an optimal scheme for uniting the reaction and separationfunctions of a process. This chapter presents many of the considerations involved in thatoptimization. Although Figure 7.1 shows only one reactor section, multiple reactor sections aresometimes required, with separation sections located between each pair of reactor sections

CD-8-28.2 LOCATING THE SEPARATION SECTION WITH RESPECT TO THE REACTOR SECTION

In many, perhaps most, chemical processes, a separation section is located after the reactionsection, as shown in Figure 7.1. In this separation section, products are purified and unconvertedreactants are recovered for recycle back to the reactor. In this manner, a process involvingreactions with unfavorable chemical equilibrium constants, Kc, at reactor conditions can achievehigh overall process conversions to desired products. Important industrial examples are thehydrogenation of nitrogen to ammonia,

N 2 + 3H 2 ↔ 2NH 3

and the hydrogenation of carbon monoxide to methanol,

CO + 2H 2 ↔ CH 3OH

both of which are exothermic reactions, whose chemical equilibrium constants, therefore,decrease with increasing temperature according to the van’t Hoff equation:

 ∂ ln K c  ∆H rxo   = (8.1)  ∂T  P RT 2

In these two examples, the chemical equilibrium constants are both less than unity andreactor conversions are less than 50% at temperatures high enough to achieve reasonable reactionrates. Because both reactions involve shrinkage in the number of moles (4 to 2 for the ammoniareaction and 3 to 1 for the methanol reaction), the reactor conversion can also be increased byincreasing the pressure, but practical considerations limit the operating pressure. However, withthe recovery and recycle of unconverted reactants, overall process conversions of 100% areapproached.

CD-8-3 Although product purification may require extreme measures to achieve productspecifications, recycle streams rarely require a significant degree of purification with respect torecycled reactants. When two or more reactants are involved, they do not have to be recoveredseparately for recycle unless their separation indexes (e.g., relative volatility) are separated by theproduct(s), as shown in the next two examples.

Example 8.1 Styrene Manufacture.

In the styrene manufacture process of Figure 10.61, the main reaction is

Methanol + Toluene → Styrene + Hydrogen + Water

The following side reaction also occurs:

Methanol + Toluene → Ethylbenzene + Water

The reactor effluent contains appreciable percentages of unreacted methanol and toluene. In this process, both styrene and ethylbenzene are products and must be purified to meet strict specifications. Water from the main reaction must be treated to the extent required for disposal to a sewer or for another use. Methanol and toluene are recovered and recycled. They are adjacent in relative volatility and, therefore, when distillation is used, they need not be separated; and because they are recycled they need not be purified to a high degree. Typically, the recycle stream might contain 5% ethylbenzene plus styrene.

Example 8.2. Cumene Manufacture.

A more complex example is the manufacture of cumene (isopropyl benzene) by the alkylation of benzene with propylene, taken from the 1997 National Student Design Competition of the AIChE. Cumene is widely used to make acetone and phenol. The fresh feeds are as follows, where the benzene feed is nearly pure, but a refinery cut of a propylene-propane mixture is used rather than a more expensive feed of nearly pure propylene.

Propylene + Benzene → n-Propylbenzene

All of the impurities in the propylene and benzene fresh feed streams, including the largeamount of propane in the propylene feed, are essentially inert, with the exception of 1-Butene, which enters into the following undesirable side reactions:

1-Butene + Benzene → t-Butylbenzene (t-BB)

1-Butene + Benzene → 1-isopropyl,4-methyl Benzene (p-Cymene)

CD-8-5Potential products and byproducts include cumene, propane, DIPBs, t-BB, p-cymene, inertlight hydrocarbons, inert aromatic compounds, and water. A main objective of the processis to maximize the production of cumene and minimize the amounts of byproduct andwaste streams. The cumene product must meet the following specifications:

Cumene purity, wt% 99.97 minimum

The propane byproduct is used as either fuel gas or LPG. Thus, it can contain water andlight hydrocarbons. However, the aromatic content cannot exceed 0.01 wt%.

Experimental alkylation data show that the two reactions above that produce DIPBs canresult in a serious loss (> 10%) of potential cumene product. To reduce this loss, tworemedies are applied, the first of which is related to Heuristic 2 in Table 5.2: (1) the use of alarge excess of benzene in the combined feed to the alkylation reactor, for example, a 4.0molar ratio of benzene to propylene to reduce the DIPB formation reactions, and (2) theaddition of a trans-alkylation reactor where the DIPBs are reacted with benzene to producecumene according to the reaction:

DIPB + Benzene → 2 Cumene

Other reactions that produce trans-alkylation heavies

Solution A preliminary block flow diagram, suggested for the cumene process, is shown inFigure 8.1. The process consists of one separation section, consisting of three columns,situated between two reactor sections, one for alkylation and one for trans-alkylation. Theseparations are all distillations, where approximate measures for the ease of distillation,

CD-8-6assuming ideal liquid solutions, are the differences between the normal boiling points of thecomponents in the alkylation reactor effluent:

Figure 8.1 Cumene process.

CD-8-8 Note that the fresh propylene feed contains approximately 31 mol%propane. Because propane is inert, Heuristic 3 of Table 5.2 should be considered.Propane can be removed in a separation section before or after the alkylationreactor. However, if removed before the reactor, a difficult separation betweenpropane and propylene is required, as discussed in Section 7.2, because the boiling-point difference is only 5.3oC (relative volatility < 1.3). In the alkylation reactor,essentially all of the propylene, as well as all of the 1-butene, are reacted.Therefore, after the reactor, propylene is not present to be separated from propane.Instead, the propane, together with water and small amounts of inert lighthydrocarbons in the propylene feed, are easily removed from the excess benzene inthe reactor effluent in the depropanizer, C1. Here, the difference in boiling pointsbetween the key components is 112.2oC (relative volatility > 10). Following thedepropanizer is a benzene-recovery distillation column, C2, where benzene isremoved, with a portion recycled to the alkylation reactor and the remainder sent tothe trans-alkylation reactor. The main separation is between benzene and cumenewith a boiling-point difference of 72.3oC (relative volatility > 5). Finally, cumeneproduct is recovered as the distillate in distillation column, C3, where the bottomsproduct, comprised of DIPBs, is sent to the trans-alkylation reactor to be convertedto cumene. In the trans-alkylation reactor, a 4.0 molar ratio of benzene to totalDIPBs is used, but the conversion of DIPBs is only 50%. By recycling the effluentfrom the trans-alkylation reactor, no net production of DIPBs is incurred. Based onlaboratory experiments and other considerations, the benzene recycle to thealkylation reactor can contain up to 10 mol% impurities. However, the combinedfeed to the alkylation reactor must not contain more than 1.3 mol% cumene.

A cardinal rule, implied in Heuristic 4 of Table 5.2, that must be adhered to

when developing a process flowsheet, is to provide exits from the process for allinert species that enter the process as impurities in the fresh feed(s) or are formed inirreversible side reactions. In the cumene process, these species include water andethane, which are more volatile than propane; isobutane, MCP, MCH, and toluene,which are more volatile than cumene; and n-propylbenzene, tBB, and p-cymene,

CD-8-9 which are more volatile than the DIPBs. Based on the product specifications for the propane and cumene products, calculations show that the total amounts of these species produced do not leave with one or both products. Consequently, two alternatives, suggested in Heuristic 4 of Table 5.2, must be evaluated. The first is to add separators to the process flowsheet. When too expensive, the second includes one or more purge or drag streams, resulting in the loss of reactant(s), product(s), or both. Two drag streams, one from the distillate of the benzene recovery column and one from the bottoms of the cumene recovery column, are used, leading to a benzene loss of about 2% and a cumene loss of less than 1%. Inclusion of drag streams and the resulting material balance calculations are the subjects of Exercise 8.1 at the end of this chapter.

Chemical processes, especially those utilizing a catalyst in the chemical reactor,

may require a feed separation section, as shown in Figure 7.1, to purify the fresh feedbefore it enters the reactor. In this separation section, catalyst poisons are removed aswell as components, other than reactants for the main reaction(s), that may enter intoundesirable side reactions in the reactor section. In general, inert chemicals can beremoved in separation sections either before or after the reactor, wherever the separationindex is more favorable, as discussed above for the cumene process. However, whenremoved after the reactor, a larger reactor is required because of the higher flow rate andlower reactant concentrations. As an example, consider the manufacture of sulfuric acid.The feed stocks are air and either sulfur or sulfide ores, where the first reaction is theoxidation of sulfur or sulfide to sulfur dioxide, the second reaction is the catalyticoxidation of SO2 to SO3, and the third reaction is the absorption of SO3 in water to formsulfuric acid. Before the first reactor, moisture must be removed from the entering air toavoid corrosion and allow the use of carbon steel. Before entering the second reactor,dust, fluorides, and arsenic and vanadium compounds must be removed from the feed gasto prevent catalyst poisoning.

What should be done when the fresh feed contains an appreciable percentage ofproduct chemicals? This occurs most frequently in isomerization reactions involving

CD-8-10light paraffin hydrocarbons, as illustrated in Example 5.2. Suppose the reaction isA ↔ B. In this case, it is important to remove the product B from the fresh feed before itenters the reactor so as to increase the rate of reaction and achieve the highest equilibriumconversion possible. However, because reactor conversion is usually incomplete forisomerization reactions, A is commonly separated from B, with A recovered andrecycled. Unless other chemicals formed in the reactor interfere with the A-B separation,the two A-B separators are combined, with the resulting separator placed before thereactor. Exercise 8.2 considers separator placement for a pentane isomerization process.

8.3 TRADEOFFS IN PROCESSES INVOLVING RECYCLE

Reactions with very large chemical equilibrium constants (e.g., > 10,000) at reactorconditions of temperature and pressure provide an opportunity for approaching 100%conversion during a single pass through the reactor. In addition, when the feed containsstoichiometric proportions of the reactants with no impurities and the reaction leads toonly one product, then in principle no separation section is needed. One such situationexists. It is the manufacture of anhydrous hydrogen chloride gas from pure, evaporatedchlorine and a stoichiometric amount of pure, electrolytic hydrogen by the reaction:

H2 + Cl2 → 2 HCl

The only pieces of equipment required are a reactor, compressors, and heat exchangers.Such a process is rare. Even when 100% reactor conversion is theoretically possible, theoptimal reactor conversion is less than 100% and a separation section is necessary. Themain reason for this is the rapid decline in reaction rate as the reacting mixture is depletedof reactants. Thus, in most processes where a chemical reactor is required, considerationmust be given to the tradeoffs between the cost of the reactor section and the cost of theseparation section that follows it.

CD-8-11 A number of factors affect the tradeoff between the reactor and separationsections, many of which were introduced in Chapters 3-7. These include

1. The fractional conversion in the reactor of the limiting reactant. This directly affects the need for and cost of the separation section.

2. The entering temperature to and mode of operation (adiabatic, isothermal,

programmed temperature profile, etc.) for the reactor. This affects heating and/or cooling costs and reactor effluent composition when side reactions are possible.

3. Reactor pressure, particularly for gas-phase reactions where the number of reactant molecules is greater than the number of product molecules. In this case, reaction kinetics may favor a higher pressure, but at the higher cost of gas compression.

4. Use of an excess of one reactant to minimize side reactions and/or increase the rate of reaction. This increases the cost of the separation system.

5. Use of an inert diluent in an adiabatic reactor to reduce the change in temperature.

This increases the cost of the separation system.

6. Use of a gas or liquid purge stream to avoid difficult separations. This reduces the cost of the separation system, but results in the loss of reactants and may increase the cost of the reactor section, depending on the purge-to-recycle ratio (ratio of purge flow rate to recycle flow rate).

The use of process simulation, in conjunction with optimization, as discussed in

Return to the toluene hydrodealkylation process in Section 4.3, with the reactionkinetics in Example 6.2. To illustrate the effect of achieving a high conversion on reactorsize, simplify the combined reactor feed by eliminating methane and neglect biphenylformation. Also, to avoid carbon formation, assume a molar ratio of hydrogen to tolueneof 5 for the combined feed to the reactor. At typical reactor conditions, the reversereaction is considered to be negligible and Eq. (6.31) gives the forward reaction rate, rf,where the Arrhenius equation for the rate constant, kf, as a function of temperature istaken from the paragraph below Eq. (6.31). Thus,

dCtoluene  −52, 000  1/2

where R = 1.987 cal/mol-K; concentrations, Ci, are in kmol/m3; time, t, is in sec; andtemperature, T, is in K. Next, the volume of both isothermal and adiabatic PFRs iscomputed for a series of conversions from 1% to 99%, for the following feed conditions:

Temperature, oF 1,200 Pressure, psia (0 pressure drop) 500

Component flow rates, lbmol/hr:

Hydrogen 2,500 Toluene 500

The calculations can be performed with any process simulator. Using theCHEMCAD program, the results for the isothermal case, plotted as reactor volumeagainst fractional conversion of toluene, are shown in Figure 8.2, with the adiabatic casein Figure 8.3. For the isothermal case, the reactor volume increases almost linearly as

CD-8-13conversion increases to 0.4. The volume then increases more rapidly until at conversionsnear 0.8, the volume turns up sharply. The reactor volume is 4,080 ft3 at a conversion of0.9, but twice that at a conversion of 0.99. Reactor Volume [1,000 ft3]

As seen in Figure 8.3, the effect of conversion on reactor volume for the adiabaticcase is very different from the isothermal case in Figure 8.2. At all conversions, thereactor volume is less for the adiabatic case. Furthermore, the difference in reactorvolumes widens as the conversion is increased. For example, at a 50% conversion, theisothermal reactor volume is 2.25 times that of the adiabatic reactor. At a 99%conversion, the ratio becomes 8. The adiabatic case benefits by the increase in

CD-8-14temperature with increasing conversion. The exothermic heat of reaction is considerableat between 21,000 and 22,000 Btu/lbmol of toluene reacted. However, the large excessof hydrogen acts as a heat carrier, curtailing the adiabatic rise in temperature.Nevertheless, the temperature increases by approximately 2.2oF per 1% increase inconversion. Thus, at 99% conversion, the reactor outlet temperature is 1,423oF. As theconversion increases, the concentration of toluene in Eq. (8.2) decreases, causing the rateof reaction to decrease. The decrease of the hydrogen concentration is not nearly aspronounced because of its large excess in the reactor feed. In the adiabatic case, thedecrease in toluene concentration with conversion is offset by the increase in the rateconstant with temperature because the activation energy is moderately high at 52,000cal/mol. This results in an approximate doubling of the rate constant with every 50oFincrease in temperature. Thus, in Figure 8.3 for the adiabatic case, unlike the isothermalcase, the increase in reactor volume is less than linear up to an inflection point at aconversion of approximately 50%. Only beyond a conversion of 90% does the reactorvolume turn up sharply.

When striving for high reactor conversions, it may be necessary to consider thereverse reaction even when the reaction is considered to be irreversible. This is the casefor the hydrodealkylation of toluene. A rate equation for the reverse reaction can bederived from the rate equation for the forward reaction, given by Eq. (8.2), by assumingthat the two rate equations are consistent with the chemical-reaction equilibrium constant.Assume that the gas reacting mixture is ideal at the high temperature of the reaction.Then, the chemical equilibrium constant can be expressed in terms of concentrations andequated to the ratio of the rate constants by:

CCH4 Cbenzene kf Kc = = (8.3) CH2 Ctoluene kb

But in chemical equilibrium, the rate of the forward reaction is equal to the rate of thebackward reaction. Therefore, from Eq. (8.2), with an as yet undetermined dependenceof component concentrations on the backward rate,

To determine the Arrhenius expression for kb from Eq. (8.3), an expression for Kc as afunction of temperature is needed. Based on the correlations of Yaws (1977), thestandard Gibbs free energy of reaction, ∆Grxo , in cal/mol, as a function of the absolutetemperature, T, in K, for the hydrodealkylation of toluene,

H2 + C7H8 → CH4 + C6H6

is given by:

∆Grxo = −11, 200 − 2.1 T (8.7)

From thermodynamics, ∆Grxo is related to the chemical-reaction equilibrium constant by

Combining Eqs. (8.6) and (8.10), the rate law for the backward reaction becomes

 −63, 200  -1/2

rb = 2.19 × 1010 exp   CH2 CCH4 Cbenzene (8.11)  RT 

When the reactor calculations are repeated for up to 99% conversion of toluene, takinginto account the reverse reaction, reactor volumes for both isothermal and adiabatic casesincrease only slightly (< 1%). This is largely due to the large concentration of hydrogen,which according to Eq. (8.11) decreases the rate of the reverse reaction. Reactionequilibrium calculations for this example give a 99.98% conversion for the isothermalcase and a 99.96% conversion for the adiabatic case. However, when only thestoichiometric quantity of hydrogen is used in the feed, the equilibrium isothermalconversion decreases to 97.3%.

8.5 RECYCLE TO EXTINCTION

In many chemical processes, the main reaction is accompanied by one or more sidereactions that produce byproducts. When the main reaction is irreversible or has a large

CD-8-17chemical-reaction equilibrium constant, but one or more of the side reactions are so-called reversible reactions with chemical-reaction equilibrium constants on the order ofone or less, the possibility of increasing the overall yield of the desired product(s) fromthe main reaction by eliminating the net production of byproduct(s) exists. This isaccomplished by applying a concept sometimes referred to as recycle to extinction. Theconcept must be applied with care and must be supported by reaction rates that aresufficiently high. This is particularly true when the main reaction is catalyzed becausethe catalyst may not support the side reaction(s). Experimental verification is essential.

The recycle to extinction concept is introduced briefly in Example 5.4 and in

Section 7.1, illustrated for the toluene-hydroalkylation process in Figure 7.4. Twoalternatives are considered: (1) production of the byproduct, and (2) recovery and recycleto extinction of the byproduct. In this process, the main reaction is the hydrogenation oftoluene to the main product, benzene, and methane:

H2 + C7H8 → CH4 + C6H6

As shown in Section 8.3, this reaction, while not completely irreversible at typical reactoroperating conditions, has a chemical-reaction equilibrium constant high enough to giveconversions greater than 99%. When the main reaction is carried out thermally, in theabsence of a catalyst, it is accompanied by the following side reaction that produces thebyproduct, biphenyl:

2 C6H6 ↔ H2 + C12H10

The chemical-reaction equilibrium constant for this reaction is written as:

C H 2 C biphenyl Kc = 2 (8.12) C benzene

Although not always considered, a further reaction to triphenyl also occurs,

CD-8-18 C6H6 + C12H10 ↔ H2 + C18H14 ,

with a chemical-reaction equilibrium constant written as:

C H 2 C triphenyl Kc = (8.13) C benzene C biphenyl

From Hougen and Watson (1947), the chemical-reaction equilibrium constant forEq. (8.12) ranges from 0.045 to 0.32 over a temperature range of 700 to 1,400oF, whilefor Eq. (8.13), the constant increases from 0.23 to 0.46 over the same temperature range.When the biphenyl and triphenyl byproducts are recovered and recycled to the reactor,they build to their equilibrium concentrations at the reactor outlet, as determined fromEqs. (8.12) and (8.13), such that no net production of either biphenyl or triphenyl occurs.In effect, the byproducts are recycled to extinction. In this manner, the production ofundesirable byproducts is eliminated and the overall yield of the main product(s) isincreased. A disadvantage of recycling the byproducts to extinction is that thebyproducts and unconverted reactants increase the cost of recycling. However, the costof the separation system downstream of the reactor may be reduced when the byproductsare recovered together with one or more of the reactants in a single recycle stream. Thisoccurs in the toluene hydrodealkylation process in which the biphenyl and triphenyl arerecovered with toluene.

A second example in which recycle to extinction should be considered is the

hydrolysis of ethylene to ethyl alcohol:

C2 H 4 + H 2 O → C 2 H5OH

which is accompanied by a reversible side reaction that produces diethylether and water,

2 C2H5OH ↔ (C2H5)2O + H2O

CD-8-19for which the chemical-reaction equilibrium constant at typical reactor conditions is 0.2.By recovering and recycling diethylether and water, the overall yield of alcohol isincreased.

A third example is the steam reforming of methane (or natural gas) in thepresence of a nickel-supported catalyst to produce synthesis gas (CO + H2), anintermediate that can be used to produce acetic acid, ammonia, gasoline, or methanol.The main reaction is:

CH4 + H2O ↔ CO + 3 H2

Typically, the reactor operation at adiabatic conditions gives an outlet temperature of

approximately 800oC, which limits the extent of the reaction to that of chemicalequilibrium, with an equilibrium constant of 126.8, with compositions in partial pressuresin atm. Reactor pressure is generally set by the available pressure of the methane andmay be as high as 30 atm.

In the presence of the catalyst, a number of side reactions occur as discussed by

Rase (1977). However, the only one of significance is the water-gas shift reaction:

CO + H2O ↔ CO2 + H2

At 800oC, the chemical-reaction equilibrium constant for this reaction is 0.929, withcompositions in partial pressures in atm. When CO2 is recovered and recycled toextinction, is the overall yield of synthesis gas increased? This is the subject of Example8.3.

Example 8.3. Steam Reforming of Naphtha.

The fresh feed to a steam reformer is 13.5 kmol/hr of methane and 86.5 kmol/hr of steam. If the outlet conditions of the reactor are 800oC and 12.2 atm and

CD-8-20chemical equilibrium is achieved for both the steam reforming and water-gas shiftreactions, determine the kmol/hr of synthesis gas produced when: (a) the CO2 produced is not recovered and recycled. (b) the CO2 is recovered from the reactor effluent and recycled to extinction.

Solution(a) At 800oC, the two chemical equilibrium equations are:

2 nCO nH3 2  P    = 126.8 nCH 4 nH2O  ntotal 

nCO2 nH 2 = 0.929 nCO nH2O

where P = 12.2 atm and ni are in kmol/hr. Since these two equations contain five unknowns, three atom-balance equations are needed. They are:

Component Fresh Feed, kmol/hr Reactor Effluent, kmol/hr

Methane 13.5 0.605

CD-8-21 From these results, 95.5% of the methane is reacted. The production of synthesis gas is 5.521 + 46.061 = 51.582 kmol/hr.

(b) For recycle of CO2 to extinction, the CO2 in the reactor effluent is recycled and added to the fresh feed to give a combined feed. At chemical equilibrium, the flow rate of CO2 in the reactor effluent is the same as that in the combined feed. The two chemical equilibrium equations remain the same, but the three atom balance equations become:

Solving the revised equations gives:

Component Combined Feed, kmol/hr Reactor Effluent, kmol/hr

Methane 13.5 0.549

Observe that there is no net production of CO2. The percent conversion of

methane is slightly greater at 95.9%, with the production of synthesis gas slightly increased to 12.946 + 38.859 = 51.805 kmol/hr. Note that in case (a), the production of CO2 from CO by the water-gas shift reaction gives an additional mole of H2 for every mole of CO2 produced. Thus, by eliminating the net production of CO2, less H2 is produced. The usual benefit of the increased yield of the main product(s) by recycle to extinction is not achieved

CD-8-22 in this case. However, in case (b), CO2 is not emitted to the atmosphere where it contributes to global warming. This is considered in more detail by Mulholland and Dyer (1999).

8.6 SNOWBALL EFFECTS IN THE CONTROL OF PROCESSES INVOLVING

RECYCLE

In recent years, chemical engineers engaged in process design in industry have becomeincreasingly aware of the need to understand the interaction of process design andprocess control when developing a control system for an entire chemical plant. When theprocess does not involve recycle, the development of the control system is relativelystraightforward because the process can be treated in a sequential manner. However, themajority of chemical processes involve recycle, for which the development of a feasibleand efficient control system, particularly for a reactor-separator-recycle network, is not atall straightforward. This is due to the possibility of the so-called snowball effect, whichrefers to a situation where a small disturbance, for example, in the fresh feed rate to areactor, causes a very large change in the flow rate of the recycle stream. When thisoccurs, either the reactor or the separation system, or both, may not be able to handle theincreased load. Whether or not the snowball effect occurs depends on the design of thecontrol system, which is the subject of Sections 20.3 (Example 20.11) and 21.5 (CaseStudy 21.3).

8.7 SUMMARY

Having studied this chapter, when designing reactor-separator-recycle networks, the

reader should

1. Understand the considerations in determining the best locations, with respect to the reactor section, of the separation sections.

CD-8-232. Be aware of the many tradeoffs between the reactor section and the separation section(s) when recycle is used.3. Know that the optimal fractional conversion of the limiting reactant in the reactor section is usually less than 100% of the equilibrium conversion.4. Be able to apply the concept of recycle to extinction to reduce waste and increase the yield of the main product.5. Be aware that the snowball effect can occur in a reactor-separator-recycle network.

Yaws, C. L., Physical Properties, McGraw-Hill, New York (1977).

EXERCISES

8.1 Cumene process with drag (purge) streams. In Section 8.2, a process for producing cumene by the alkylation of benzene with propylene is described. The flowsheet for the process is given in Figure 8.1. However, that flowsheet does not provide for the removal of water, ethane, isobutane, MCP, MCH, toluene, n-propylbenzene, tBB, and p-cymene. For their removal, it is proposed to add two drag (purge) streams to the flowsheet: one from the distillate of the benzene recovery column,

CD-8-24 C2; the other from the bottoms of the cumene recovery column, C3. Also, the flowsheet in Figure 8.1 does not provide for an exit for the heavies produced in the alkylation and trans-alkylation reactors in the event that their amounts are too large to be included in the allowable impurity in the cumene product. Thus, it may be necessary to add a fourth distillation column, C4, following C3, with the distillate from C4 fed to the trans-alkylation reactor and the bottoms from C4 being a heavies product. If so, the heavies must not contain more than 5% of the DIPBs and lighter entering C4.

Most of the data for the cumene process is given in Section 8.1. However, missing are the product distributions for the two reactors. These are as follows from laboratory studies:

Trans-alkylation Reactor Alkylation Reactor Change in pounds per 100 Component Change in pounds per 100 pounds of propylene in the pounds of propylene in the combined feed to the combined feed Alkylation Reactor

CD-8-25 Note, again, that the conversion of DIPBs in the trans-alkylation reactor is only 50%.

Using the above data and that in Section 8.1, revise the flowsheet in Figure 8.1 and produce a complete material balance with the component flow rates in lbmol/hr for each stream in your flowsheet. Try to maximize the production of cumene. Be sure to add two drag streams for removal of byproducts, and a fourth distillation column, if necessary. Compute the overall percent conversion of benzene to cumene and the annual production of cumene in lb/yr if the operating factor is 0.95. If a heavies product is produced, what could it be used for?

8.2 The feed to a pentane isomerization process consists of 650 kmol/hr of n-pentane and 300 kmol/hr of isopentane. The effluent from the catalytic isomerization reactor will contain 6.5 moles of isopentane for every mole of n-pentane. The catalyst prevents the formation of neopentane. If the isopentane product, produced by separating isopentane from n-pentane by distillation, is to contain only 2 wt% n- pentane and the separation system is to be placed before the reactor, calculate the total flow rate and composition of the reactor effluent, the combined feed to the reactor, and the bottoms product from the distillation column. Design the distillation column. Repeat the material balance calculations and the design of the distillation column if the separation system is placed after the reactor. Based on your results and without determining any capital or operating costs, which separation system placement is preferred?

CD-8-26 Chapter 9

Second-Law Analysis

9.0 OBJECTIVES

The first law of thermodynamics is widely used in design to make energy balances aroundequipment. Much less used are the entropy balances based on the second law of thermodynamics.Although the first law can determine energy transfer requirements in the form of heat and shaftwork for specified changes to streams or batches of materials, it cannot even give a clue as towhether energy is being used efficiently. As shown in this chapter, calculations with the secondlaw or a combined first and second law can determine energy efficiency. The calculations aredifficult to do by hand, but are readily carried out with a process simulation program. When thesecond-law efficiency of a process is found to be low, a better process should be sought. Theaverage second-law efficiency for chemical plants is in the range of only 20-25%. Therefore,chemical engineers need to spend more effort in improving energy efficiency.

After studying this chapter, the reader should

1. Understand the limitations of the first law of thermodynamics.

2. Understand the usefulness of the second law and a combined statement of the first and second laws. 3. Be able to specify a system and surroundings for conducting a second-law analysis. 4. Be able to derive and apply a combined statement of the first and second laws for the determination of lost work or exergy. 5. Be able to determine the second-law efficiency of a process and pinpoint the major areas of inefficiency (lost work). 6. Understand the causes of lost work and how to remedy them. 7. Be able to use a process simulation program to perform a second-law analysis.

CD-9-19.1 INTRODUCTION

A chemical process uses physical and/or chemical operations to transform feed materialsinto products of different composition. Table 9.1 lists the types of operations that are most widelyused. Depending on the production rate and the operations used, the process is conductedbatchwise, continuously, or cyclically. A continuous, heat-integrated process that illustratesseveral of the operations in Table 9.1 is shown in Figure 9.1, where benzene and a mixture ofxylene isomers are produced by the disproportionation of toluene. The heart of the process is afixed-bed catalytic reactor, R-1, where the main chemical change is the reaction 2C7H8 → C6H6 + C8H10 isomers

This reaction is conducted in the presence of hydrogen to minimize the undesirable formation ofcoke by condensation reactions. However, other undesirable side reactions such as C7H8 + H2 → C6H6 + CH4occur and produce light paraffins. Chemicals in the reactor effluent are separated from each otheras follows. Hydrogen is recovered for recycle by partial condensation in exchanger E-2 with phaseseparation in flash drum D-1; light paraffin gases are removed in fractionator C-1; benzene isrecovered and purified in fractionator C-2; and mixed xylenes are recovered and purified, andunreacted toluene is recovered for recycle in fractionator C-3. Compressors K-1 and K-2 bring

CD-9-2Figure 9.1 Process for disproportionation of toluene to benzene and xylenes.

CD-9-3fresh hydrogen and recycled hydrogen, respectively, to reactor pressure. Pump P-1 brings freshtoluene to reactor pressure. Pumps P-2, P-3, and P-4 deliver reflux to fractionators C-1, C-2, andC-3, respectively. Pumps P-3 and P-6 deliver benzene and xylene products, respectively, tostorage, and pump P-5 recycles toluene. Furnace F-1 uses the combustion of fuel oil with air tobring reactants to reactor temperature, after preheater E-1 has recovered a portion of the thermalenergy in the reactor effluent. Cooling water is used in overhead condensers E-4, E-6, and E-9, andsteam is used in reboilers E-5, E-7, and E-10 of fractionators C-1, C-2, and C-3, respectively.Benzene and xylene products are cooled by water in coolers E-8 and E-11 (not shown in Figure9.1) before being sent to storage. Exchanger E-3 preheats feed to fractionator C-1 with bottomsfrom the same fractionator. Cooling water is supplied mainly by recycle from cooling tower T-1by pump P-7. Electricity for all pumps and compressors, and steam for reboilers is produced fromcoal-fired power plant B-1. The overall input to and output from the process is representedschematically in Figure 9.2.

Figure 9.2 Overall process streams for toluene disproportionation.

Ideally, each operation in a process would be conducted in a reversible manner to achieve

the minimum energy input or the maximum energy output, corresponding to a second-lawthermodynamic efficiency of 100%. Even if this were technically feasible, such a process would beuneconomical because of excessive capital investment in equipment, which would have to beessentially infinite in size to minimize transport gradients. Nevertheless, it is economical to modifyexisting processes to reduce energy consumption, and to design new processes to operate at higher

CD-9-4thermodynamic efficiencies. A second-law thermodynamic analysis identifies inefficient processesand the operations within these processes that are the most wasteful of energy, so that the processengineer can direct his or her efforts to conserving energy.

9.2 THE SYSTEM AND THE SURROUNDINGS

To conduct a second-law analysis, a process is divided into a system and surroundings. The systemis the matter contained in the operating unit(s) on which the engineer wishes to focus. Everythingnot in the system is in the surroundings. The boundaries of the system may be real or imaginary,rigid or movable, and open or closed to the transfer of matter between the system and thesurroundings. Some references call a closed system simply a system, and an open system, intoand/or out of which matter can flow, a control volume. They refer to the boundary of the controlvolume as the control surface across which matter can flow.

Batch, cyclic, and continuous processes are shown schematically in Figure 9.3. Batch andcyclic processes are usually divided into a closed system (or simply a system) and surroundings;continuous processes are divided into an open system (or control volume) and surroundings.

Figure 9.3 Common methods of processing.

The division of a process into system and surroundings is the choice of the one performingthe thermodynamic analysis. Many choices are possible for a chemical process. For example, inFigure 9.1, the system can be the complete process, with the surroundings being the ambient air,water, and so forth, surrounding the equipment (commonly referred to as the infinite surroundings,dead state, or infinite heat reservoir) and the storage tanks for the raw materials and products.

More commonly, utility plants (e.g., the steam power plant and cooling-water system) areconsidered separately from the rest of the process. This is shown schematically in Figure 9.4,where the process is divided into three systems. The benzene-mixed xylenes plant is sufficientlycomplex that it is advisable to divide it into a reaction section and a separation section, as shown inFigure 9.5. Any individual operation in the process - for example, fractionator C-2 - can be thesystem and everything else the surroundings. Finally, a portion of a single operation can be thesystem - for example, one tray in fractionator C-2.

9.3 ENERGY TRANSFER

Heat or work, or both, can be transferred across the boundaries of closed or open systems. If noheat is transferred across its boundaries, the system is said to be adiabatic or thermally isolated; andif neither work nor heat is transferred, the system is said to be totally isolated.

The most useful kind of energy transfer is work. For example, a rotating or reciprocatingshaft at the boundary of a system causes shaft work. Less useful, but more common, is heattransfer, which occurs when the temperatures of the system and the surroundings differ. If thesystem is at the higher temperature, it loses energy and the surroundings gain energy; and if thesystem is at the lower temperature, it gains energy and the surroundings lose energy.

A number of devices are used in processes to transfer work between a system and itssurroundings. Pumps, compressors, blowers, and fans convert shaft work into fluid energy for themain purpose of increasing fluid pressure. Turbines and expanders take energy from a fluid,causing fluid pressure to decrease, and convert the energy to shaft work for use elsewhere. Amotor converts electrical work to shaft work. A generator converts shaft work to electrical work.

As an example of energy transfer by work, consider Figure 9.6(a), where an incompressible

liquid at 25oC having a specific volume, V, of 0.001 m3/kg is pumped continuously at a rate m of

CD-9-710 kg/s from a pressure P1 of 0.1 MPa to a pressure P2 of 2.0 MPa, with no change in kinetic orpotential energy, by a rotating shaft driven by an electrical motor. In the absence of electricalresistance, shaft friction, and fluid friction, Electrical work input to the electric motor = shaft work delivered to the pump by the motor = shaft work delivered to the liquid by the pump = isothermal, isokinetic, isopotential energy increase of liquid ( P2 − P1 ) = 10(0.001)(2,000,000 - 100,000) = W = mV

= 19 kN-m/s (kJ/s or kW)

In actual equipment (as shown in Figure 9.6(b)), electrical resistance may permit only a95% transfer of electrical work to the motor shaft, shaft friction may permit only a 90% transfer ofshaft work to the fluid, and fluid friction may cause a rise in fluid temperature equivalent to a 5%loss of the shaft work. For the same increase in fluid pressure, the electrical work input to theelectric motor is then

19 Winput = = 23.39 kW (0.95)(0.90)(0.95)

CD-9-8 The difference, 23.39 - 19.00 = 4.39 kW, between the rate of electrical work input to themotor and the rate of energy required to increase the fluid pressure is the power not used inaccomplishing the desired goal. This excess power causes temperatures in the system and/or thesurroundings to rise.

If the temperature of a system or a part of the surroundings remains reasonably constant

when heat transfer between these two regions occurs, then the system or the part of thesurroundings is called a heat reservoir. Heat reservoirs include heating media, such as steam, hotwater, Dowtherm, oil, molten salts, mercury, and flue gases produced by combustion; and coolingmedia such as air, water, chilled water, ammonia, propane, and other refrigerants. For each ofthese reservoirs, it is convenient to assign a temperature. It is also convenient to distinguishbetween finite-sized heat reservoirs, which are designed to operate at certain desired temperatures,Ti, and the essentially infinite heat reservoirs that exist in the natural environment, such asatmospheric air, oceans, and large lakes or rivers at temperatures designated as T0.

9.4 THERMODYNAMIC PROPERTIES

When work and/or heat is transferred to or from a system, energy changes occur. The mostcommon forms of energy are those associated with (1) macroscopic motion (kinetic energy), (2)location in a gravitational field (potential energy), and (3) internal energy due to translational,rotational, and vibrational motions of molecules, atoms, and electrons; together with the potentialenergy due to forces acting between molecules, atoms, electrons, and nuclei. The first two forms ofenergy are taken relative to some arbitrary reference, such as a point on the surface of the earth. Inmost chemical processes, changes to these two forms of energy are relatively small and are oftenignored. An exception is the combustion chamber and nozzle of a rocket engine, where the heat ofreaction (internal energy) is converted to kinetic energy. Internal energy is most important inchemical processing and is taken relative to some arbitrary reference condition.

CD-9-9 The internal energy of a substance is a state property, because its value depends on the stateor condition of the substance, which is determined by temperature, pressure, composition, phase (ifmore than one phase is possible), and the reference condition. Changes in internal energy areindependent of the path employed in moving from one state to another.

Another state property, closely related to internal energy, is enthalpy, defined by therelation H = U + PV (9.1)This property is particularly convenient for continuous processes because the two terms on theright-hand side frequently appear together in energy balance equations.

The most desirable reference conditions for internal energy and enthalpy in processes wherechemical reactions take place are 0 K or 25oC, zero pressure, and standard chemical elements, suchas C (graphite), H2 (gas), O2 (gas), N2 (gas), Cl2 (gas), and S (rhombic sulfur), rather than thechemical species themselves that are in the mixture. With this reference condition, internal energyand enthalpy changes automatically take into account heat of reaction. Felder and Rousseau (2000)discuss this reference condition. As an example, the enthalpy of 1 kg of superheated steam at300oC and 1 MPa relative to the elements H2 (gas) and O2 (gas) at 0 K and 0 Pa is determined to be-12,209.3 kJ. Alternatively, from the steam tables in van Wylen et al. (1994), for a referencecondition of saturated liquid water at 0oC, the enthalpy is 3,051.2 kJ/kg.

It is well known from thermodynamic principles that energy transferred as work is moreuseful than energy transferred as heat. Work can be completely converted to heat, but only afraction of heat can be converted to work. Furthermore, as the temperature of a system isdecreased, heat transferred from the system becomes less useful and less of the heat can beconverted to work. A state property that accounts for the differences between heat and work isentropy, S. When heat is transferred into a closed system at temperature T, the entropy of thesystem increases because entropy transfer accompanies heat transfer. By contrast, work transfer(shaft work) is not accompanied by entropy transfer. When heat is transferred at a rate Q from asurrounding heat reservoir at a constant temperature, Treservoir, into a system, the heat reservoirexperiences a decrease in entropy given by

CD-9-10 −Q ∆Sreservoir = (9.2) Treservoirwhere ∆S is the entropy change in Btu/hr-oR. The lower the value of T, the greater the decrease inentropy.

For a pure, ideal gas, only temperature affects U and H. However, the entropy, S, of anideal gas is affected by both temperature and pressure. Accordingly, the reference pressure for Uand H is usually taken as zero. For S, the reference pressure is usually taken as 1 atm to avoid avalue of S equal to minus infinity. At a reference temperature of 0 K, the entropy of a crystallinesubstance is zero, by the third law of thermodynamics.

Figure 9.8 Isothermal flow through a pipe

Ideal Gas Mixing. When C species are mixed at constant pressure and temperature, asillustrated in Figure 9.9, the change in the entropy flow rate is given by Eq. (9.10), appliedseparately for each species j: P m j ∆S1→ 2 j = m j R ln (9.11) Pjwhere Pj = xjP is the partial pressure of species j, and xj is its mole fraction, and m j is its molar

Figure 9.9 Isothermal mixing of C ideal gas species.

Thermodynamic Availability

When matter is taken from state 1, at a given velocity, elevation, composition, temperature T, andpressure P, to state 2, at a different velocity, elevation, composition, T, and P, it is of interest todetermine the maximum amount of useful work that can be extracted or the minimum amount ofwork that is needed. Ignoring kinetic energy and potential energy differences and referringenthalpies to the elements, the first law of thermodynamics can be used to determine the netamount of energy transferred by heat and/or work in moving from state 1 to state 2, which issimply the change in enthalpy. The first law cannot be used to determine the maximum orminimum amount of useful work, which depends on the details of the process used to effect thechange in state. The maximum or minimum is achieved only if the process is reversible.

To determine the maximum rate at which work is performed, Wmax , in bringing a stream toequilibrium with its surroundings, a reversible path can be selected, as illustrated in Figure 9.10. Astream at molar flow rate, m , in state 1, at T1 and elevated pressure, P1, is fed to turbine I, whichoperates adiabatically and reversibly. It is expanded to P2 and the environmental temperature, T0,while producing shaft work at the rate, W s I . The effluent stream from turbine I is expanded

isothermally (non-adiabatically) and reversibly in turbine II to the environmental pressure, P0. Thepath is shown in the P-V and T-S diagrams, the second of which shows the isentropic behavior ofturbine I.

CD-9-13 WsI

WsII m , T1 , P1 I T0 , P2

II Q I = 0 T0 , P0

Q II P T1 P1

P2 T0 P0 S1 S2 V

Figure 9.10 Reversible path.

Ignoring kinetic and potential energy changes, the first law of thermodynamics, applied tothe overall process is: m ∆H1→0 = Q − W s (9.14)where Q is the rate of heat transfer to turbine 2, and W s is the sum of the shaft work ratesdelivered by the two turbines.

Substituting in the first law, Eq. (9.14):

This reversible work is the maximum work “available” in bringing the feed stream to theenvironmental conditions; that is, W s is the maximum rate of obtaining work, which can be written

m A1→0 . The intensive property, A1→0 , was initially referred to as the thermodynamic availabilityand is commonly referred to as the exergy. The concept of availability was first developed in detailby Keenan (1951).

It follows that the change in availability of a stream, when it is converted from state 1 tostate 2 in a chemical process, as shown in Figure 9.11, is: ∆A1→2 = A2 − A1 = ∆H1→2 − T0 ∆S1→ 2 (9.19)That is, the change in the maximum work available from the stream is a function solely of itschanges in enthalpy and entropy, and the environmental temperature. Like H and S, A is a statefunction, independent of path, but dependent on the temperature, T0, and pressure, P0, of the deadstate. If chemical reactions occur, the availability also depends on the composition of the deadstate.

T1, P1 T2, P2 A1 A2

Figure 9.11 Availability change upon processing.

Typical Availability Changes

In this subsection, availability changes are computed for several simple processes to show thesignificant impact of the change in entropy. These are taken from the monograph by Sussman(1980), who presents many other excellent examples, including three that take into accountchemical reaction, one of which deals with a complete methane reforming process. In all cases, theenvironmental (dead-state) temperature in the following examples is taken as 298 K = 537oR.

CD-9-15 Superheating Steam. As shown in Figure 9.12, saturated steam at 250 psia and 401°F issuperheated isobarically to 600°F, with the enthalpy and entropy values taken from the steamtables. Substituting in Eq. (9.19):

H1 = 1,201.1 Btu/lb H2 = 1,319 Btu/lb

Although the enthalpy of the stream is increased by 117.9 Btu/lb, which equals the heattransferred to the stream, the maximum work that can be obtained from stream 2, if it is taken tothe environmental conditions, is increased by only 51.4 Btu/lb, which is less than 50 % of the heattransferred, because the entropy term increases so significantly.

Liquefying Air. As shown in Figure 9.13, air at 25°C and 1 atm is condensed isobaricallyto a saturated liquid at -194.5°C. Substituting in Eq. (9.19):

∆A1→ 2 = ∆H1→2 − T0 ∆S1→ 2

CD-9-16Note that the enthalpy and entropy data are obtained from the air tables, where the reference state issaturated liquid air at 25°C. The change in enthalpy, -101.37 kcal/kg, is the heat removed from thecondenser, using a refrigerator that requires considerable compression work. In this case, theentropy change is sufficiently negative to cause the entropy to be about three times more positivethan the negative enthalpy change. This causes a large increase in the availability of the liquid air.Stated differently, 174.6 kcal/kg is the maximum work obtained from the liquid air in returning it tothe environmental state, and is the minimum work of refrigeration in liquefying air.

Air Sat’d. Liquid

1 2 25°C -194.5°C 1 atm Condenser 1 atm

H1 =127.11 kcal/kg H2 = 25.74 kcal/kg

S1 = 0.9260 kcal/kg-K S2 = 0

Figure 9.13 Condensation of air.

Throttling. As shown in Figure 9.14, superheated steam is throttled adiabatically across a

When throttling, the entire change in availability is due to the negative change in entropy. Stateddifferently, the entropy term is the maximum loss of the ability of the stream to do work intransferring to its environmental (dead) state. Using Eq. (9.19), A1 is computed to be 434.9 Btu/lb,and consequently, 12% of its “available” work is lost in throttling. As considered subsequently in

CD-9-17this chapter, the possibility of replacing the valve with a turbine to recover power should beconsidered when the pressure of a stream must be reduced.

= 0 + 298(0.79 ln0.79 + 0.21 ln0.21) 1.987

cal = −304.3 mol airThe positive entropy change upon mixing results in the negative change in availability. Stateddifferently, 304.3 cal of work are the minimum required to separate air into nitrogen and oxygengases.

Figure 9.16 Thermal mixing of water.

The availability change upon thermal mixing is illustrated conveniently in an availability

flow diagram, as shown in Figure 9.17, where the widths of the arrows are approximatelyproportional to the availability flow rates. Combining the availability flow rates for the hot andcold streams, the availability flow rate entering the mixer is 4.05 + 0.555 = 4.61 kcal/s. In themixer, this is divided into 0.99 kcal/s, which leaves in the mixed effluent stream, and 3.62 kcal/s,which is lost to the environment; that is, approximately 78% of the availability to do work is lostupon thermal mixing. Clearly, this loss decreases as the temperatures of the hot and cold streamsapproach each other. Sussman (1980) makes extensive use of availability flow diagrams like thatin Figure 9.17.

CD-9-19 0.99 kcal/s Hot 4.05 kcal/s Mixer

Cold 0.555 kcal/s 3.62 kcal/s

Figure 9.17 Availability flow diagram for thermal mixing of water.

9.5 EQUATIONS FOR SECOND-LAW ANALYSIS

In this section, the first and second laws of thermodynamics are used to derive useful equations forcomputing the lost work of any process. A general energy balance (first law of thermodynamics)can be written for a system bounded by the control volume shown in Figure 9.18. Streams atcertain fixed states flow at fixed rates into or out of the control volume, heat and work aretransferred at fixed rates across the boundaries of the control volume, matter within the controlvolume undergoes changes in amount and state, and the boundaries of the control volume expandor contract. The energy balance for such a control volume over a period of time, ∆t, is ∆ ( mU )sys )flowing streams = Q 0 − ∑ Q i − ∑ Wi + ∆ ( mH (9.20) ∆t i i

where ∆ ( mU )sys is the change in internal energy of the system, ∆ ( mH )flowing streams is the sum of

enthalpy flows leaving the system minus the sum of those entering the system, Q 0 is positive for

heat transfer from the infinite surroundings at T0 to the control volume, and Q i is positive for heattransfer to the control volume from a heat reservoir at temperature Ti different from T0. Eq. (9.20)ignores changes in kinetic energy and potential energy for both the system and the flowing streams.The term ∑ W i is positive for work done by the system on the surroundings and includes i

An entropy balance for the system in Figure 9.18 can be written in a manner analogous tothat used for the energy balance, Eq. (9.20), except that here we prefer to write an entropy balancefor both the control volume and the surroundings. The result is ∆ ( mS )sys Q 0 Q + ∆ ( mS )flowing streams − − ∑ i = ∆Sirr (9.21) ∆t T0 i Ti

where ∆ ( mS )sys is the change of entropy of the system, ∆ ( mS

)flowing streams is the sum of entropy

flows leaving the system minus the sum of those entering the system, − Q 0 T0 is the rate ofdecrease in entropy of the infinite surroundings when heat is transferred from the infinitesurroundings at T0 to the system in the control volume, and −∑ Q i Ti ( ) is the sum of the rates of

entropy decrease in the various heat reservoirs at various temperatures, Ti, that are used to transferheat into the system. Unlike energy, entropy is not conserved. The term ∆Sirr is the increase inentropy of the universe due to the process. It is zero only for a reversible process. Otherwise, it ispositive and is a measure of the irreversibility of the process.

Although ∆Sirr is a fundamental quantity, it is of limited practical use because of the

difficulty in interpreting the significance of its magnitude. As with another fundamentalthermodynamic quantity, chemical potential, it is preferred by chemical engineers to use a

CD-9-21surrogate property. For chemical potential, that quantity is fugacity; for ∆Sirr , it is availability(exergy), which was defined earlier and arises naturally, as will be shown next, when the first andsecond laws of thermodynamics are combined.

To derive availability, combine Eqs. (9.20) and (9.21) by eliminating Q 0 . The result is

∆  m (U − T0 S )  sys  T  + ∆  m ( H − T0 S )  flowing streams − ∑ Q i 1 − 0  + ∑ Wi + T0 ∆Sirr = 0 (9.22) ∆t i  Ti  iIn this equation, in the second term on the left-hand side, we see that the enthalpy and entropyappear together to form a combined factor that is similar to the Gibbs free energy. However, theentropy is multiplied by the dead-state temperature, T0, instead of the stream temperature, T. Inaddition, the first term on the left side can be rewritten to give the same combination, H − T0 S , bysubstituting Eq. (9.1), the definition of enthalpy, for the internal energy. The result is ∆  m ( H − T0 S − PV )  sys  T  + ∆  m ( H − T0 S )  flowing streams − ∑ Q i  1 − 0  + ∑ Wi + T0 ∆Sirr = 0 (9.23) ∆t i  Ti  i We now define an availability function, B, for the combination of enthalpy and entropy inEq. (9.23): B = H – T0S (9.24)The availability function in Eq. (9.24) and availability in Eq. (9.19) differ from each other in thatthe availability is referenced to a dead state at T0, P0, and a composition for every element in theperiodic table) and is, therefore, an absolute quantity. The availability function, by contrast, can bereferenced to any state and is not an absolute quantity. In Eq. (9.23), however, only the change inavailability function appears. By their definitions, the change in availability function is exactlyequal to the change in availability. When evaluating a process, only the change in availability oravailability function, ∆A or ∆B, respectively, is important. If one is interested in the maximumuseful work that can be extracted from a material that is brought to equilibrium with the dead state,then the availability, A, is of importance. In the second-law analysis of a process, we will use ∆B.

CD-9-22 In addition, we also note in Eq. (9.22) that ∆Sirr is multiplied by T0 and that their product

has the units of energy flow. Accordingly, it is given the name lost work, LW , or loss ofavailability or exergy: LW = T0 ∆Sirr (9.25) Substitution of Eqs. (9.24) and (9.25) into Eq. (9.23) gives

Alternatively, Eq. (9.26) may be rearranged to the following form:

For a reversible process, ∆Sirr and, therefore, T0 ∆Sirr and LW , are zero. For an irreversible

process, ∆Sirr and LW are positive. The lost work represents the energy flow (power) lost because

of irreversibilities in the process. The lost work is much easier to relate to than ∆Sirr .

The significance of Eq. (9.27) is best illustrated by a simple case. Consider a continuous,steady-state, adiabatic process, where Eq. (9.27) simplifies to

∑W + LW = −∆  m ( B )

i i flowing streams (9.28)

If the process decreases the availability function for the flowing streams, then the right-handside of Eq. (9.28) will be a positive quantity. That decrease will be converted to useful work doneon the surroundings and/or lost work. However, if the lost work is greater than the decrease inavailability, work will have to be transferred from the surroundings to the processing system. If theprocess is also reversible, then ∑W i i is the maximum work that can be extracted from the decrease

in availability. Thus, for such a reversible process,

   ∑ Wi  = −∆  m ( B )  flowing streams , for ∆B = (−) (9.29)  i  max If the process increases the availability function for the flowing streams, then the right-handside of Eq. (9.28) will be a negative quantity. That increase will require work to be done by the

CD-9-23surroundings on the process (i.e., a negative value for ∑W i i ). If lost work (a positive quantity)

occurs in the process because of irreversibilities, then, according to Eq. (9.28), an equivalentamount of additional work must be done on the process by the surroundings to satisfy the change inavailability function. If the process is reversible, then ∑W i i is the minimum work required for the

increase in availability. Thus, for such a reversible process,

   ∑ Wi  = −∆  m ( B )  flowing streams , for ∆B = (+ ) (9.30)  i  min Eqs. (9.26) and (9.27) are availability balances. The heat and the work terms are transfers ofavailability to or from the process. For a continuous, steady-state process, let us compare anenergy balance to an availability balance. The comparison is facilitated by rewriting Eq. (9.20) forthe energy balance and Eq. (9.26) for the availability balance, respectively, in the following forms,where work and heat terms are all positive because they are labeled into or out of the system:

1. The left-hand side of Eq. (9.31) is zero. That is, energy is conserved. The left-hand side of Eq. (9.32) is zero only for a reversible process. Otherwise, the left-hand side is positive and availability is not conserved. In an irreversible process, some availability is lost.

CD-9-24 2. In the energy balance, work and heat are counted the same. In the availability balance, work and heat are not counted the same. All work input increases the availability of material flowing through the process. Only a portion of heat transferred into a system is available to increase the availability of flowing streams. The heat is degraded by a coefficient equal to 1 − (T0 / T ) . This coefficient is precisely the Carnot

cycle efficiency for a heat engine that takes heat from a source at temperature, T, and converts a portion of it to useful work, discharging the balance to a sink at a lower temperature, T0. Note that in the availability balance, T is not the temperature of the process stream within the system, but is the temperature of the heat source or sink outside the system. 3. The energy balance, which is valid whether the process is reversible or not, has no terms that take into account irreversibility. Thus, the energy balance cannot be used to compute the minimum or maximum energy requirements when taking material from inlet to outlet states. The availability balance does have a term, LW , that is a measure of irreversibility. When the lost work is zero, the process is reversible and Eq. (9.32) can be used to determine the maximum or minimum energy requirements to cause a change in availability.

Regardless of whether a net availability of heat or work is transferred to or from a process,

the energy balance must be satisfied. Thus, the energy and availability balances are used together todetermine energy requirements and irreversibilities that lead to lost work. The more efficient aprocess, the smaller the lost work.

9.6 EXAMPLES OF LOST-WORK CALCULATIONS

Before proceeding with a discussion of the second-law thermodynamic efficiency in the nextsection, two examples are provided to illustrate the calculation of lost work for chemical processes.

CD-9-25EXAMPLE 9.1For the first example, consider the continuous two-stage compression of nitrogen gas shown inFigure 9.19, which is based on actual plant operating conditions. The system or control volume isselected to exclude the electric power generation plant and cooling-water heat sink. Assume thatthe temperature, T, of the cooling water is essentially equal to the dead-state temperature, T0.Calculate the lost work.

The flow rate of nitrogen through the process is 3,600 lb/hr. Therefore, the change inavailability of nitrogen is 3,600[-663.24 - (-744.24)] = 291,600 Btu/hr

Because the availability increases, energy must be transferred into the system. The electrical powerinput of 107.3 kW is equivalent to 366,400 Btu/hr. This is greater than the availability increase,which represents the minimum energy input corresponding to a reversible process. Thus, thecompression process has irreversibility. To determine the extent of the irreversibility, substitute thechange in availability of the nitrogen, and the work input into the availability balance, Eq. (9.33),for the lost work:

This is equivalent to 21.9 kW or 29.4 hp.

Where does the irreversibility occur? To answer this, separate second-law analyses areneeded for each of the two compressors and the intercooler. Unfortunately, data on the nitrogenleaving the first compressor and leaving the intercooler are not provided. Therefore, these separateanalyses cannot be made.

How can we apply the first law of thermodynamics to the nitrogen compression problem?We can apply an energy balance to calculate how much heat must be transferred from the nitrogento cooling water in the intercooler: Welectrical in = Q out + m ( H 2 − H1 ) (9.34)Therefore,

CD-9-27Note that the enthalpy increase of 180,100 Btu/hr is far less than the minimum amount of energy of291,600 Btu/hr that must be added. Even if the compressors and the intercooler were reversible,291,600 - 180,100 = 111,500 Btu/hr of energy would have to be transferred out of the system.Although this is considerably less than the 186,300 Btu/hr for the actual process, it is still a largeamount.

EXAMPLE 9.2As a second example, consider the plant operating data shown in Figure 9.20 for a propanerefrigeration cycle. Saturated propane vapor (state 1) at 0 oF and 38.37 psia for a flow rate of 5,400lb/hr is compressed to superheated vapor (state 2) at 187 psia and 113oF. The propane is thencondensed with cooling water at 77oF in the refrigerant condenser to state 3 at 98.7oF and 185 psia.Reducing the pressure across the valve to 40 psia causes the propane to become partially vaporized(state 4) at the corresponding saturation temperature of 2oF. The cycle is completed by passing thepropane through the refrigerant evaporator, where the propane absorbs heat from the matter beingrefrigerated and from which it emerges as a saturated vapor (state 1), thus completing the cycle.Calculate the lost work.

Figure 9.20 Operating conditions for propane refrigeration cycle.

CD-9-28SOLUTIONLet the system be circulating propane and the electric motor drive of the compressor, but not thecooling water used in the condenser or the matter being refrigerated in the evaporator.

For each pass through the cycle, there is no net enthalpy change for the propane. Theenergy balance, if applied incorrectly, would therefore indicate that no energy is required to run thecycle. But, of course, energy input is required at the compressor, and heat is transferred to thesystem from the matter being refrigerated at the evaporator. By an energy balance, the sum ofthese two energy inputs is transferred out of the system to cooling water at the condenser.

Again, it is emphasized that the first law of thermodynamics cannot be used to determineminimum or maximum energy transfer to or from a system. Instead, we must use the second law orthe availability balance (combined first and second laws). For the propane refrigeration cycle, theavailability balance of Eq. (9.32) simplifies to  T0   T0  LW = Win +  1 −  T  Q in −  1 −  Qout (9.35)  Evaporator   TCondenser  For a reversible cycle, the lost work would be zero, and this form of the availability balanceis the classical result for the refrigeration (reverse Carnot) cycle. To prove this, the first law gives Win + Q in = Q out (9.36)

Substitution of this equation into the lost-work equation, Eq. (9.35), with LW = 0 , so as toeliminate Q out , gives the following widely used equation for the coefficient of performance (COP)of a refrigeration cycle: Q in TEvaporator COP = = (9.37) Win TCondenser − TEvaporator

The lost work for the cycle is computed in the following manner. First, we take the dead-state temperature, T0, to be the cooling-water temperature, TCondenser. The lost work then reduces to  T0  LW = Win + 1 −  T  Q in (9.38)  Evaporator  The electrical work input is given in Figure 9.20 as 70 kW. The heat transferred in theevaporator is obtained most readily from an energy balance on the propane as it flows from state 3

In a reversible cycle, with LW = 0, only 70 - 45.05 or 24.95 kW of electrical work input would berequired.

9.7 THERMODYNAMIC EFFICIENCY

The thermodynamic efficiency of an operation or an entire process depends on its main goal andthe work lost in accomplishing that goal. Goals differ from application to application. For

CD-9-30example, the main goal of an adiabatic turbine operating continuously might be to produce work.The main goal of a refrigeration cycle might be the transfer of heat from the stream beingrefrigerated to the refrigerant. In continuous chemical processes that involve reactors, separators,heat exchangers, and shaft-work devices, the main goal is the increase or decrease of theavailability function of the streams flowing across the boundaries of the system. For a complexbatch chemical process, the main goal is the increase or decrease of the batch availability functionm(B - PV) of the system.

To derive general expressions for thermodynamic efficiency, we write Eq. (9.26), thecombined energy and entropy balance, in the form

Each of the terms on the right-hand side of Eq. (9.40) represents a possible main goal. Theavailabilities of some main goals are listed in Table 9.2. The thermodynamic efficiency iscomputed from one of two equations, depending on the sign of the term that represents the maingoal on the right-hand side of Eq. (9.40). If the sign is positive, the thermodynamic efficiency isgiven by main goal − LW η( + )goal = (9.41) main goal

CD-9-31If the numerical value of the main goal selected is negative, the thermodynamic efficiency is givenby main goal η( − )goal = (9.42) main goal − LW The application of Eq. (9.42) always results in a positive efficiency that is equal to or lessthan unity (i.e., 100%), because the main goal has a negative sign and the lost work is greater thanor equal to zero. On the other hand, Eq. (9.41) can give values ranging from less than zero up tounity. A negative efficiency results when the lost work is greater than the absolute value of themain goal. For example, consider a continuous process in which the main goal is to decrease theavailability function of the flowing streams. If the process is reversible and exchanges heat onlywith the infinite surroundings, then LW = 0 and work could be done on the surroundings. If,however, the process is so irreversible that, instead, work must be done on the system by thesurroundings, then LW will be greater than the main goal, −∆ ( mB )flowing streams , and Eq. (9.41) will

The application of Eqs. (9.41) and (9.42) for the calculation of thermodynamic efficiencymay be illustrated by considering the two examples in the preceding section. For the continuous,steady-state, steady-flow, two-stage compression process shown in Figure 9.19, the main goal is tochange the availability function of the nitrogen gas. The calculations previously presented give

Main goal = − m ( B2 − B1 ) = −291,600 Btu/hr

LW = 74,800 Btu/hr

Because the main goal has a negative value, we apply Eq. (9.42) to obtain

−291, 600 η= = 0.796 or 79.6% −291, 600 − 74,800

This is consistent with the previous calculation of 20.4% for the loss of input electrical energy.

CD-9-32 In the refrigeration cycle of Figure 9.20, the main goal - the transfer of heat from the matterbeing refrigerated at 10oF to the propane refrigerant - requires the work of a reversible Carnotcycle, which was calculated to be -25 kW. This work was accompanied by 45 kW of lost work.Thus, Eq. (9.42) gives −25 η= = 0.357 or 35.7% −25 − 45

9.8 CAUSES OF LOST WORK

Lost work is caused by irreversibilities; their major causes are:

1. Mixing of two or more streams or batches of material that differ in temperature,

pressure, and/or composition. Such mixing leads to significant increases in entropy, but may be unavoidable when preparing a composite feed for chemical reaction. Often, however, such mixing can be avoided when recycling material. Quenching a hot stream with a cold stream increases entropy.

2. Finite driving forces for transport processes. For reasonable-size processing equipment, finite driving forces are needed for heat transfer and mass transfer. However, the smaller the driving forces, the smaller is the lost work. In distillation, small driving forces are best achieved with countercurrent flow of vapor and liquid at reflux ratios close to minimum. For heat exchangers, small temperature-driving forces are achieved with countercurrent flow and small temperature approaches at either end of the exchanger.

3. Fluid friction and drag. Significant decreases in skin friction for flow of fluids in pipes can be achieved by increasing pipe diameter, thereby reducing fluid velocity. Reducing the velocity or streamlining the shape of the object can reduce form drag for flow of fluid past submerged objects.

CD-9-33 4. Chemical reactions occurring far from equilibrium. To minimize lost work, reactions should be carried out with little or no dilution, with minimal side reactions, and at maximum yields to avoid separations and byproduct formation. This is best achieved by using selective catalysts. If the reaction is exothermic, it is best carried out at high temperature to maximize the usefulness of the energy produced. If the reaction is endothermic, it is best carried out at below ambient temperature to utilize heat from the dead state.

5. Transferring heat to cooling water, especially when that heat is available at an elevated temperature. Good uses should be found for waste heat.

6. Mechanical friction in machinery such as pumps, compressors, and turbines.

Second-law thermodynamic efficiency of the majority of chemical processes is in the range

of 25 - 30%. Economic analyses have shown that it is worthwhile to seek ways to improve thisefficiency to at least 60%. Machinery is available with efficiencies of 80% and higher.

9.9 THREE EXAMPLES OF SECOND-LAW ANALYSIS

In this section, three detailed examples of second-law analysis are presented for chemicalprocesses. Each example includes the calculation of lost work, the determination of where the lostwork occurs, and consideration of how the lost work can be reduced. The examples involve (1) thepropane refrigeration cycle introduced in Section 9.6, (2) the separation of a mixture of propyleneand propane by distillation, and (3) a process for the hydrogenation of benzene to cyclohexane.The third example is computed with ASPEN PLUS.

EXAMPLE 9.3 A Refrigeration Cycle

In Sections 9.6 and 9.7, the total rate of lost work and overall thermodynamic efficiency of apropane refrigeration cycle, shown in Figure 9.20, is calculated. Now, consider this cycle in detail

CD-9-34to determine where the lost work occurs with respect to each of the four steps in the cycle. Then,attempt to improve the efficiency of the cycle by concentrating on those steps where most of thelost work occurs. Although the overall process is a cycle, each separate step in the cycle can betreated as a continuous process so that Eq. (9.27) applies.

Because state 4 is a partially vaporized condition, the fractions of vapor and liquid must bedetermined to obtain S4. That is, if ψ is the weight fraction vaporized, then S4 = ψ ( S4 )V + (1 − ψ )( S4 ) L

Table 9.3 Lost Work for Propane Refrigeration Cycle

Step in Cycle State to State LW ( kW ) Percentage of Total

How can the thermodynamic efficiency of this refrigeration cycle by improved? Table 9.3shows that the major loss is due to the compressor, with moderate losses in the refrigerantcondenser and the valve, but only a small loss in the refrigerant evaporator. Some improvementscan be made by maintaining the same basic cycle, but adjusting the operating conditions andchanging the equipment to accomplish the following:

1. Increase the efficiency of the compressor.

2. Reduce the frictional pressure drop in the refrigerant condenser. Use a higher-temperature coolant for the refrigerant condenser or reduce the compressor discharge pressure to lower the temperature of the refrigerant at states 2 and 3.

CD-9-37 3. Replace the valve with a power-recovery turbine.

4. Reduce the frictional pressure drop in the refrigerant evaporator. Increase the pressure at state 4 to reduce the temperature-driving force in the refrigerant evaporator.

Figure 9.21 Revised propane refrigeration cycle.

A revised cycle that incorporates these improvements is shown in Figure 9.21. Comparisonof the cycle with the original one in Figure 9.20 shows the following:

1. The valve is replaced by a power-recovery turbine that supplies a portion of the power required by the compressor.

CD-9-38 4. The compressor inlet and discharge pressures are changed from 38.37 psia to 44.35 psia and from 187 psia to 154.9 psia, respectively, thus reducing the compression ratio from 4.874 to 3.493. The corresponding changes in refrigerant temperature cause reductions in the minimum temperature-driving forces in the condenser and evaporator from 21.7oF (98.7 - 77) to 8oF (85 - 77) and from 8oF (10 - 2) to 2oF (10 - 8), respectively.

Next, the lost work is calculated assuming that the power-recovery turbine and thecompressor operate isentropically. Also, the rate of heat transfer in the refrigerant evaporator isassumed to be the same as for the original cycle (597,200 Btu/hr, as calculated above). Requiredthermodynamic properties of propane for the revised cycle are

It is worthwhile to begin calculations with the refrigerant condenser, where the known heatduty permits us to determine the propane flow rate.

From State 4 to State 1

Q i = m ( H1 − H 4 )

Therefore, Q i 597, 200 m = = H1 − H 4 −684.6 − H 4

To obtain H4, note that since the power-recovery turbine is assumed to operate isentropically, S4 =S3 = 1.0805 Btu/lb-oR. Also note that (S4)V > 1.0805 > (S4)L. Therefore, state 4 is partiallyvaporized propane. If ψ is the weight fraction vaporized,

From State 3 to State 4

From State 1 to State 2

Letting −WC = rate of work transferred by the compressor to the propane,

−WC = m ( H 2 − H1 )

The enthalpy, H2, depends on the temperature of the propane leaving the compressor. It can beobtained by noting that S2 = S1 = 1.3486 Btu/lb-oRbecause of the isentropic compression assumption. From the thermodynamic data given,

(S ) > (S ) > (S ) 100o F V 2 V 90o F V

By interpolation, T2 = 96.32 °F and H2 = -659.9 Btu/lb. Therefore,

WC = −4, 766  −659.9 − ( −684.6 )  = −117, 720 Btu/hr

CD-9-40Of this amount, 18,110 Btu/hr is supplied from the power-recovery turbine. Therefore, thetheoretical electrical power input, W E , is

EXAMPLE 9.4 Separation of a Propylene-Propane Mixture by Distillation

The initial design of a distillation operation for the continuous, steady-state, steady-flowseparation of a propylene-propane mixture is shown in Figure 9.22. Conventional distillation isused with a bottoms pressure of 300 psia so that cooling water can be used in the partial condenserto provide reflux. The relative volatility of propylene to propane is quite low, varying from 1.08 to1.14 for conditions at the top of the fractionator to conditions at the bottom of the fractionator,respectively; thus, a large external reflux ratio of 15.9 is required at operation near the minimumreflux. Because of high product purities, as well as the low average relative volatility, 200 stagesare required at 100% tray efficiency. With 24-in. tray spacing, two columns in series are neededbecause a single column would be too tall. Therefore, an intercolumn pump is shown in addition tothe reflux pump. Total pressure drop for the two columns is 20 psi.

As shown in Figure 9.22, the system is chosen so that it does not include the 77oF coolingwater used as the coolant in the partial condenser or the 220oF saturated steam used as the heating

CD-9-41medium in the partial reboiler. Enthalpies of the feed stream and the two product streams are givenin Table 9.4, with reference to the elements H2 (gas) and C (graphite) at 0oR and 0 psia using theSoave-Redlich-Kwong (SRK) equation of state with standard heats of formation. Entropies givenare referred to 0oR and 1 atm.

The thermodynamic efficiency is computed from Eq. (9.41) because the main goal is tochange the availability function of the streams, which is −∆ ( mB )flowing streams = −140.81 kW

Thus, as in most continuous separation operations, the availability function of the flowing streamshas been increased. In this example, the increase is brought about mainly by the transfer of heat inthe reboiler, giving −∆ ( mB ) −140.81 η= = = 0.0689 or 6.89% −∆ ( mB ) − LW −140.81 − 1,902.58

This is a very low efficiency, but typical of conventional distillation of mixtures with a low relativevolatility because of the large energy expenditures required in the reboiler. Therefore, otherseparation methods, such as adsorption, have been explored for this application. Also, elaborateschemes for reducing the reboiler heat duty in distillation have been devised, including multieffectdistillation and operation at lower pressures using heat pumps, as discussed in Section 10.9. Onesuch alternative scheme, using reboiler-liquid flashing, for the separation of propylene frompropane is shown in Figure 9.23. The feed is reduced in pressure to 108 psia by a power-recoveryturbine and then distilled in a single column operating at a bottoms pressure of 112 psia. Liquidleaving the bottom tray is flashed across an expansion valve to a pressure corresponding to asaturation temperature lower than the saturation temperature of the overhead vapor so that thepartial condenser can be used as a reboiler. A compressor is needed to return the reboiled vapor tothe bottom of the column. Because the required reboiler duty is somewhat larger than the requiredcondenser duty, an auxiliary steam-heated reboiler is needed. The large reduction in reboiler steamis somewhat offset by the power requirement of the compressor.

Although the lost work is much lower than the value of 1,902.58 kW computed for the system inFigure 9.22, the thermodynamic efficiency is still low. The two cases are not really comparablebecause the product conditions are not the same.

EXAMPLE 9.5 A Process for Converting Benzene to Cyclohexane

Here, a process is considered that involves a chemical reactor as well as separators, heatexchangers, and pumps. A continuous, steady-state, steady-flow process for manufacturingapproximately 10 million gallons per year of high-purity cyclohexane by the catalytichydrogenation of high-purity benzene, at elevated temperature and pressure, is shown in Figure9.24. The heart of the process is a reactor in which liquid benzene from storage, together withmakeup hydrogen and recycle hydrogen in stoichiometric excess, take part in the reaction

CD-9-45 Figure 9.24 includes all major equipment and streams together with a set of operatingconditions for making a preliminary design and second-law analysis. As shown, 92.14 lbmol/hr ofpure liquid benzene feed (S1) at 100oF and 15 psia is pumped by P1 to 335 psia and mixed in-lineand adiabatically at M1 with impure hydrogen makeup gas (S3) containing 0.296 mol% nitrogen at120oF and 335 psia, gas recycle (S4), and a cyclohexane recycle (S5) to produce the combinedreactor feed (S6). In the cooled reactor, R1, 99.86% of the benzene in stream S6 is hydrogenated toproduce the saturated-vapor reactor effluent (S7) at 392oF and 315 psia. This effluent is reduced intemperature to 120oF at 300 psia by the cooler, H1, and then separated at these conditions in thehigh-pressure flash drum, F1, into a hydrogen-rich vapor and a cyclohexane-rich liquid. A total of8.166% of the vapor from this flash drum is purged to stream S11 at line tee D1, with theremaining vapor (S12) recycled to the reactor, R1, to provide an excess of hydrogen. At the linetee, D2, 62% of the liquid (S10) from flash drum F1 is sent in stream S14 to a low-pressureadiabatic flash drum, F2, at 15 psia. Gas from F2 is vented to stream S15, while liquid is taken ascyclohexane product S16. The remaining liquid S13 from F1 is recycled by pump P2 to reactor R1to control the pressure of the saturated-vapor reactor effluent.

It is convenient to use computer simulation to perform mass and energy balance

calculations automatically for continuous-flow, steady-state processes like the one in Figure 9.24.For this example, ASPEN PLUS is used. This requires that the process flow diagram be convertedto a simulation flowsheet as discussed in Section 4.3. That flowsheet is shown in Figure 9.25, inwhich each stream has a unique name, the same as or similar to that shown in Figure 9.24. Eachoperation is a simulation unit within which two names appear. The top name, e.g., R1 for thereactor, is a unique user-specified unit name or so-called block i.d. The bottom name, e.g.,RSTOIC for the reactor, refers to the selected ASPEN PLUS model, or subroutine, for theoperation. As discussed earlier, in many cases, a particular operation can be simulated with two ormore models. The information given in Figures 9.24 and 9.25 is sufficient to prepare the input fora simulation. As discussed earlier, specifications can be entered interactively in the ASPEN PLUSprogram. Specifications entered on input forms are converted by ASPEN PLUS to a compactlisting that can be displayed if desired. The listing is given in Figure 9.26, where the flowsheettopology is followed by the list of components with user-selected names followed by data banknames. Thermodynamic properties are computed by option SYSOP1, which is the Chao-Seader

CD-9-46method with the Grayson-Streed constants for estimating K values and the Redlich-Kwongequation of state for obtaining the departure functions for the effect of pressure on enthalpy andentropy. All mixture enthalpies and entropies are referenced to the elements at 25oC. Therefore,energy and entropy balances automatically account for enthalpy and entropy changes due tochemical reaction. This greatly simplifies the calculations when chemical reactions occur as in thiscyclohexane process. The availability function, B, is readily computed from its definition, Eq.(9.24), for a selected value of T0. Specifications for the two inlet streams, S1 and S3, follow. TheASPEN PLUS program concludes with the operating conditions for each simulation unit.

Figure 9.25 ASPEN PLUS flowsheet for the cyclohexane process.

In Figure 9.25, two recycle loops are clearly seen. However, no recycle convergencemethod is specified in the ASPEN PLUS program, and the flowsheet does not show theconvergence units. Accordingly, ASPEN PLUS selects, by default, the tear streams, initialcomponent flow rates of zero for the tear streams, and a convergence method. The convergedresults of the simulation for the ASPEN PLUS program in Figure 9.26 are given in Figure 9.27,where component and total molar flow rates, temperature, pressure, molar enthalpy, molar vaporand liquid fractions, molar entropy, density, and average molecular weight are listed. Bycomparing streams S1 and S16, it is seen that the overall yield of cyclohexane from the process is91.2899/92.1400 or 99.08%. By comparing streams S3 and S16, it is seen that an overall excess of

CD-9-47[(282.9599/3)/91.2899] - 1.0 or only 3.32% H2 is used. Examination of stream S4 or S12 showsthat relatively little N2 is recycled, although the amount is large relative to the N 2 in the makeuphydrogen. The amount of cyclohexane recycle in stream S5 or S13 is considerable compared to thebenzene feed S1. The energy balance results are summarized in Table 9.5, where the net energytransfer rates are listed for each operation, and are considerable for the reactor, R1, and the partialcondenser, H1.

Figure 9.26 ASPEN PLUS input in paragraph form for the cyclohexane process.

CD-9-50 The results in Figure 9.27 and Table 9.5 are used to perform a second-law analysis. Thedead-state temperature is taken as 100oF. The calculation of lost work for the entire process andthe corresponding second-law efficiency is carried out conveniently on a spreadsheet bytransferring results from ASPEN PLUS, as shown in Figure 9.28. Note that the availabilityfunction for each stream can be computed and printed by ASPEN PLUS. The overall efficiency isonly 25.7%. Similar analyses are carried out readily for the separate operations in the process. Thefraction of the total lost work for each operation is as follows:

Operation % of Total Lost Work

This table shows clearly that the reactor and cooler are, by far, the largest contributors to theinefficiency of the process. Some reduction in lost work can be achieved by replacing the partialcondenser with two or three heat exchangers operating with coolants at different temperaturelevels. But what can be done with the reactor? Would it be better to operate it at a lower or highertemperature? Should a larger excess of hydrogen be used? Clearly, there is room for considerableimprovement in the reactor operation. See Exercise 9.23.

CD-9-519.10 SUMMARY

Having studied this chapter, the reader should

1. Know the differences between and the limitations of the first and second laws of thermodynamics.

2. Understand the concepts of the irreversible change in entropy and lost work or exergy.

3. Be able to use a process simulator to compute lost work and second-law efficiency.

4. Be able to pinpoint major causes of lost work in a process and determine ways to improve the efficient use of energy.

9.1 A stream of hot gases at 1,000oC, having a specific heat of 6.9 cal/mol-oC, is used to preheat air fed to a furnace. Because of insufficient insulation, the hot gas cools to 700oC before it enters the air preheater. How much availability per mole does it lose?

9.2 An ideal gas, with Cp = 7 cal/mol-oC, is compressed from 1 to 50 atm while its temperature rises from 25 to 150 oC. How much does its availability change per mole?

9.3 Superheated steam at 250 psia and 500oF is compressed to 350 psia. The isentropic efficiency of the compressor is 70%. For the compressor, compute its a. Lost work b. Thermodynamic efficiency

9.4 Steam at 400oF, 70 psia, and 100 lb/hr is compressed to 200 psia. The electrical work is 4.1 kW. Determine the a. Lost work b. Thermodynamic efficiency c. Isentropic efficiency

9.5 The rate of heat transfer between Reservoir A at 200oF and Reservoir B at 180oF is 1,000 Btu/hr. a. Compute the lost work. b. Adjust the temperature of Reservoir A to 10oF. For the same heat duty and lost work, compute the temperature of Reservoir B. How do the approach temperatures, ∆TAB, compare?

9.6 Nitrogen gas at 25oC and 1 atm, with Cp = 7 cal/mol-K, is cooled to -100oC at 1 atm. Assuming an ideal gas, calculate the minimum work per mole required for cooling. What is the maximum work per mole that can be obtained when the gas is returned to 25oC and 1 atm?

CD-9-539.7 An equimolar stream of benzene and toluene at 1,000 lbmol/hr and 100oF is mixed with a toluene stream at 402.3 lbmol/hr and 50oF, as discussed in connection with Figures 4.8 and 4.9. Assuming ideal vapor and liquid mixtures, use a process simulator to compute the a. Change of availability upon mixing b. Lost work c. Thermodynamic efficiency

9.8 Consider the cooler, H2, in the monochlorobenzene separation process in Figure 4.23 and 4.24. Assume that the heat is transferred to an infinite reservoir of cooling water at 77 oF. a. Using the enthalpy and entropy values in the results for the sample problem in the ASPEN PLUS section of the CD-ROM that accompanies this textbook, determine the lost work associated with the cooler. b. Let the reservoir be at 100oF and repeat (a).

9.9 Two streams, each containing 0.5 lb/hr steam at 550 psia, are mixed as shown:

a. Compute the heat loss to an environmental reservoir at 77oF.

b. Compute the lost work and thermodynamic efficiency.

9.10 1,000 lb/hr of saturated water at 600 psia is superheated to 650oF and expanded across a turbine to 200 psia, as illustrated.

CD-9-54 Calculate the a. Isentropic efficiency of the turbine b. Lost work for the process c. Thermodynamic efficiency of the process

9.11 Superheated steam at 580 oF and 500 psia is expanded across a turbine, as shown below, to 540oF and 400 psia. 0.9 kW of shaft work are produced. The turbine exhaust is cooled by a 77oF reservoir to its dew point at 400 psia.

Determine the a. Flow rate of steam in lb/hr b. Isentropic efficiency of the turbine c. Lost work d. Thermodynamic efficiency

9.12 Calculate the minimum rate of work in watts for the gaseous separation at ambient conditions indicated in the following diagram.

CD-9-559.13 Calculate the minimum rate of work in watts for the gaseous separation at ambient conditions of the feed indicated below into the three products shown.

CD-9-589.17 A light-hydrocarbon mixture is to be separated by distillation, as shown in Figure 9.29, into ethane-rich and propane-rich fractions. Based on the specifications given and use of the Soave-Redlich-Kwong equation for thermodynamic properties, use ASPEN PLUS with the RADFRAC distillation model to simulate the column operation. Using the results of the simulation, with T0 = 80oF, a condenser refrigerant temperature of 0oF, and a reboiler steam temperature of 250oF, calculate the a. Irreversible production of entropy, Btu/hr-oR b. Change in availability function in Btu/hr c. Lost work in Btu/hr, kW, and Hp d. Thermodynamic efficiency

Figure 9.29 Distillation process for Exercise 9.17.

CD-9-599.18 A mixture of three hydrocarbons is to be separated into three nearly pure products by thermally coupled distillation at 1 atm, as shown in Figure 9.30.

Figure 9.30 Thermally coupled distillation process for Exercise 9.18.

Based on the specifications given and other specifications of your choice to achieve reasonably good separations, together with use of the Peng-Robinson equation for thermodynamic properties, use ASPEN PLUS with the MULTIFRAC distillation model to simulate the column. Using the results of the simulation, with T0 = 100oF, calculate the a. Irreversible production of entropy, Btu/hr-oR b. Change in availability function in Btu/hr c. Lost work in Btu/hr, kW, and Hp d. Thermodynamic efficiency

CD-9-609.19 Consider the hypothetical perfect separation of a mixture of ethylene and ethane into pure products by distillation as shown in Figure 9.31.

Two schemes are to be considered: conventional distillation and distillation using a heat pump with reboiler liquid flashing. In both cases the column will operate at a pressure of 200 psia, at which the average relative volatility is 1.55. A reflux ratio of 1.10 times minimum, as computed from the Underwood equation, is to be used. Other conditions for the scheme using reboiler liquid flashing are shown below. Calculate for each scheme: a. Change in availability function (T0 = 100oF) b. Lost work c. Thermodynamic efficiency Other thermodynamic data are Latent Heat of Vaporization (Btu/lbmol) Ethylene at 200 psia 4,348 Ethane at 200 psia 4,751 Ethane at 90 psia 5.473

CD-9-619.20 Consider a steam engine that operates in a Rankine cycle, as illustrated below:

The turbine exhaust is a saturated vapor.

a. Find the saturation temperature of the turbine exhaust. b. For an isentropic efficiency of 90 percent, determine the shaft work delivered by the turbine. What is the temperature of the feed to the turbine? c. Compute the lost work for the turbine. d. Compute the thermodynamic efficiency for the turbine.

9.21 A reactor is to be designed for the oxidation of sulfur dioxide, with excess oxygen from air, to sulfur trioxide. The entering feed, at 550 K and 1.1 bar, consists of 0.219 kmol/s of nitrogen, 0.058 kmol/s of oxygen, and 0.028 kmol/s of sulfur dioxide. The fractional conversion of sulfur dioxide is 50%. The reaction is very exothermic. Three cases are to be considered: 1. Adiabatic reaction. 2. Isothermal reaction with the heat of reaction transferred to boiler feed water at 100oC. 3. Isothermal reaction with the heat of reaction transferred to boiler feed water at 200oC.

For each case, compute the lost work in kW.

9.22 For the revised propane refrigeration cycle in Figure 9.21 (Example 9.3), let the isentropic efficiencies of the turbine and compressor be 0.9 and 0.7, respectively. Compute the a. Lost work for the four process units and the entire cycle. b. Thermodynamic efficiency of the cycle.

CD-9-629.23 Alter the design of the cyclohexane process in Example 9.5 to reduce the lost work and increase the thermodynamic efficiency. Use a simulation program to complete the material and energy balances, and compute the entropies and availability functions for all of the streams, as well as the lost work for each piece of equipment.

9.24 The chilled-water plant at the University of Pennsylvania sends chilled water to the buildings at 42°F and receives warmed water at 55°F. A refrigerant is vaporized in the refrigerant condenser at 38°F, as it removes heat from warmed water. The refrigerant is condensed to a saturated liquid at 98°F. The condensing medium is water at 85°F, which is heated to 95°F as it absorbs heat rejected from the refrigerant. The warmed condenser water is cooled in a cooling tower, in which it is sprayed over a stream containing ambient air. Assume that the ambient air is at 100°F and 95% humidity on a hot summer day and is rejected at 100% humidity. For Phase I of the plant, the cooling capacity is 20,000 tons. a. Calculate the flow rates of the chilled water and condenser water in gal/min. b. Select a refrigerant and its operating pressures. Assuming an isentropic efficiency of 70% for the compressor, determine the refrigerant flow rate and the brake horsepower for the compressor. c. Calculate the lost work and thermodynamic efficiency.

9.25 Consider the solar or waste-heat refrigeration cycle in Figure 9.32, which was proposed by Sommerfeld (2001). In addition to the conventional refrigeration loop, a portion of the condensate is pumped to an elevated pressure, where it is vaporized using solar energy or low-temperature waste energy in a chemical complex. Its saturated vapor effluent is expanded to recover power in a turbine and mixed with the gases from the compressor.

CD-9-63 Figure 9.32 Solar or waste-heat refrigeration cycle.

Use a process simulator to solve the material and energy balances for the followingspecifications: R-134a refrigerant 4-ton refrigeration load at 20°F Refrigerant evaporator effluent - saturated vapor at 40°F Condenser heat rejected to environment at 77°F Condenser effluent - saturated liquid at 125°F Solar or waste-heat available at 220°F Solar or waste-heat collector effluent - saturated vapor at 200°F Isentropic efficiency of the compressor = 70% Isentropic efficiency of the turbine = 90% Isentropic efficiency of the pump = 100%a. Determine the flow rates of refrigerant in both loops; the three operating pressures; the condenser and collector heat duties; the power consumed or generated by the compressor, pump, and turbine; the coefficient of performance, lost work, and thermodynamic efficiency for the refrigerator.b. Vary the condenser effluent temperature to determine its effect on the solution in part a.

ASPEN IPE FOLDERS AND FILES 51

APPENDIX I - DEPROPANIZER – ASPEN PLUS REPORT 52

APPENDIX II - DESIGN CRITERIA SPECIFICATIONS 58

APPENDIX III - ASPEN IPE CAPITAL ESTIMATE REPORT FOR THE 61

DEPROPANIZER

APPENDIX IV - ASPEN IPE CAPITAL ESTIMATE REPORT FOR THE 70

MONOCHLOROBENZENE SEPARATION PROCESS

CD-IPE-i INTRODUCTION

These notes are prepared to provide a step-by-step procedure for estimation of thetotal capital investment using the Aspen Icarus Process Evaluator (Aspen IPE). AspenIPE is a software system provided by Aspen Technology, Inc., for economic evaluationof process designs. It determines the capital expenditure, operating costs, and theprofitability of proposed designs. Aspen IPE has an automatic, electronic expert systemwhich links to process simulation programs. It is used to (1) extend the results of processsimulation, (2) generate rigorous size and cost estimates for processing equipment, (3)perform preliminary mechanical designs, and (4) estimate purchase and installation costs,indirect costs, the total capital investment, the engineering-procurement-constructionplanning schedule, and profitability analyses.

Aspen IPE usually begins with the results of a simulation from one of the majorprocess simulators (e.g., ASPEN PLUS, HYSYS, CHEMCAD, and PRO/II), it beingnoted that users can, alternatively, provide equipment specifications and requestinvestment analysis without using the process simulators. In these notes, only resultsfrom ASPEN PLUS are used to initiate Aspen IPE evaluations and only capital costestimation is emphasized. Readers should refer to the Aspen IPE User’s Guide (press theHelp button in Aspen IPE) for detailed instructions, explanations, and for improvementsin new versions of the software system.

These notes are organized as follows:

1. Instructions are provided to prepare an ASPEN PLUS simulation for use with Aspen IPE.

2. A depropanizer example is provided to illustrate the use of Aspen IPE. The

depropanizer is a distillation tower to recover propane and lighter species from a normal-paraffins stream, as shown in Figure 1. The simulation flowsheet and selected results are shown in Appendix I and in the multimedia tutorial on the CD-ROM that contains these course notes (ASPEN → Tutorials → Separation Principles → Flash and Distillation). Also, a copy of the file, RADFRAC.bkp, is provided on the CD-ROM.

3. Additional features of Aspen IPE are introduced for a more complete process, the monochlorobenzene (MCB) separation process, which is discussed in Sections 4.4 of the textbook (Seider et al., 2004). A copy of the simulation file, MCB.bkp, is provided on the CD-ROM that contains these notes.

After completing these notes, to practice estimating capital costs using AspenIPE, you may wish to solve Exercises 16.4 and 17.21 in the textbook.

In these notes, all of the calculations were carried out using Aspen IPE, Version11.1, with the design and cost basis date being the First Quarter 2000.

CD-IPE-1 Figure 1 Depropanizer

PREPARING AN ASPEN PLUS SIMULATION FOR ASPEN IPE

To estimate equipment sizes and costs using Aspen IPE for a process simulatedwith ASPEN PLUS, it is necessary to prepare the simulation results for use with AspenIPE. While this is accomplished in a similar manner for most of the major processsimulators, these notes focus on the steps to prepare ASPEN PLUS simulations. For thesteps when using the other process simulators, the reader should refer to the Aspen IPEUser’s Guide (press the Help button in Aspen IPE).

It is normally necessary to adapt the simulation file in two ways. First, to

estimate equipment sizes, Aspen IPE usually requires estimates of mixture properties notneeded for the material and energy balance, and phase equilibria calculations performedby the process simulators. For this reason, it is necessary to augment the simulationreport files with estimates of mixture properties, such as viscosity, thermal conductivity,

CD-IPE-2and surface tension, for the streams in the simulation flowsheet. Second, Aspen IPErequires specifications to estimate equipment sizes that are not computed by some of theapproximate simulation models. This is the case, for example, when the DISTL andRSTOIC models are used in ASPEN PLUS. These must be replaced by more rigorousmodels, such as the RADFRAC and RPLUG models. This replacement can be viewed asthe first step in computing equipment sizes and costs. Note that it is also possible toprovide specifications for computing equipment sizes without using ASPEN PLUS.

Additional Mixture Properties

Estimates for the additional stream properties are added using the PROPSETS.aptfile on the CD-ROM that contains these course notes. To accomplish this, the ASPENPLUS simulation file is opened first; e.g., RADFRAC.bkp (which is available on the CD-ROM that contains these course notes). Under the File pull-down menu, the Import entryand the PROPSETS.apt file are selected. Aspen IPE automatically adds three newproperty sets, after which the file can be saved as RADFRAC-IPE.bkp, a copy of whichis on the CD-ROM that contains these notes. To check that this has been accomplished,using the Data pull-down menu, select Setup and then Report Options. Then, display theStreams page by selecting the appropriate tab and click the Property Sets button.Observe that all three Aspen IPE property sets have been entered into the SelectedProperty Sets box. Now that the Aspen IPE property sets have been added, it isnecessary to re-run the simulation.

INVESTMENT ANALYSIS USING ASPEN IPE

In this section, the use of Aspen IPE for equipment sizing and costing isillustrated for a depropanizer and for the monochlorobenzene separation process.

DEPROPANIZER

This example involves the single distillation column shown in Figure 1, with itssimulation flowsheet and selected results shown in Appendix I and on the multimediatutorial on the CD-ROM that contains these course notes (ASPEN → Tutorials →Separation Principles → Flash and Distillation).

Initial Setup

Having sent the ASPEN PLUS simulation file to Aspen IPE, it is openedautomatically and the Create New Project dialog box appears:

CD-IPE-3The user can either select an existing project in which to start a new scenario, or enter anew Project Name. The Project Name RADFRAC-IPE is assigned automatically fromthe ASPEN PLUS file name, however punctuation marks are not allowed, so enter theProject Name DEC3 instead. Note that the underscore and space characters arepermitted. After pressing the OK button, the first of four dialog boxes, not shown here,appear. The first is the Project Properties dialog box, in which a Project Description andfurther remarks may be entered. A units of measure set is also chosen, which for thisexample is the Inch-Pound (IP) units set.

Second, the Input Units of Measure Specifications dialog box is displayed. Thisform allows the user to customize the units of measure that will appear on inputspecification forms. Click the Close button to accept the default settings.

Third, the General Project Data dialog box appears. Since no adjustments areneeded in this example, press the OK button. Fourth, the Load Simulator Data? dialogbox is displayed. Enter Yes to do so.

Aspen IPE now opens two windows shown below. The narror Project Explorer,on the left, is in Project View mode, and a wider Main window, initially blank, is on theright. Note that two additional windows, Palette and Property, can be opened using theView pulldown menu. Aspen IPE allows the user to specify many parameters forequipment sizing or to accept default values. These are the basis for sizing the equipmentand for specifying its utilities. The first step in completing this simulation is to examinethe project Design Criteria. This can be done by selecting the Project Basis View tab inthe Project Explorer. Note that the Design Criteria and Utility Specifications entriesunder the Process Design heading are the most relevant when estimating equipment sizesand costs. Double-click on Design Criteria to cause the Design Criteria-IP form toappear in the Main window:

CD-IPE-4Default values are provided for many of the entries, but they can be modified asnecessary, and missing entries can be entered. Particular attention should be paid to thedesign pressure and temperature, to the overdesign factors, to the residence times in theprocess vessels, as well as to other tower information. The user must be careful to checkall of the relevant specifications that apply to the equipment under study. Note that thedesign criteria are defined in the Aspen IPE User’s Guide, which can be accessed usingthe Help button in Aspen IPE, with the values specified for the depropanizer processshown in Appendix II (Defining the Project Basis → Process Design → DesignCriteria). Note also that design criteria files can be created for use with other designprojects. For implementation details, see the Aspen IPE User’s Guide.

Also, it is usually important to examine the default values associated with theutilities. For this purpose, the Utility Specifications entry under the Process Designheading is selected to produce the Develop Utility Specifications dialog box:

CD-IPE-5Note that all existing utilities to be used by Aspen IPE are listed. Default values shouldbe examined and modified, and missing utilities should be added. For example, becausethe textbook recommends that process designs accept cooling water at 90°F and heat it to120°F, it is necessary to replace the temperatures associated with the cooling waterutility. To modify these temperatures, double-click on the Cooling Water entry, whichproduces the Utility Specifications dialog box:

Then, the inlet and exit temperatures are changed to 90 and 120°F. Other default valuescan be changed similarly. Click OK when finished.

CD-IPE-6 To add a utility not in the existing utility list, click on the Create option on theDevelop Utility Specifications dialog box. As shown below, low-pressure steam is addedas a utility, which is named Steam @50PSI and has the Steam Fluid Class.

After the Create button is pressed, the new utility is displayed as shown below, where theentries have already been made from the steam tables of Smith et al. (2001).

When complete, the OK button is pressed to return to the Develop Utility Specificationdialog box. Then, the Close button is pressed to return to the IPE Main window. Notethat utility files can be created for use on other design projects. For implementationdetails, see the Aspen IPE User’s Guide.

CD-IPE-7 Other specifications can be changed in a manner similar to those described for theutilities and design criteria. More information and definitions are provided inthe Aspen IPE User’s Guide (Defining the Project Basis → Process Design).

Mapping Process Simulation Units into Aspen IPE

Having completed the initial setup, the next step is to map the process simulation units(that is, blocks, modules, or subroutines) into more descriptive models of process equipment(e.g., mapping a HEATX simulation unit into a floating-head, shell-and-tube heat exchanger;mapping a RADFRAC simulation unit into a tray tower, condenser, reflux accumulator, etc.) andassociated plant bulks, which include installation items, such as piping, instrumentation,insulation, paint, etc. After Aspen IPE completes the mapping and reserves storage for theinstallation items, equipment sizes are computed. Note that the mapping and equipment sizingsteps are accomplished in sequence, with sizes and costs of the installation items estimatedduring the Equipment Costing step. To begin the mapping step in the IPE Main window, theMap Simulator Items button on the toolbar is pressed to produce the Map dialog box:

For the depropanizer, all items are mapped and sized in sequence, since the SizeICARUS Project Components button is checked. When this button is not checked, onlythe mapping step is completed. Also, when there are multiple process units of a certaintype, it may be preferable to map each process unit independently. For example, if twodistillation towers differ in tray efficiency, it is necessary to map them separately andchange the tray efficiency under Design Criteria before each tower is mapped. In thiscase, with just one tower, it is simplest to press the Map all Items button under Source.Under Basis, the Default and Simulator Data button should be selected, as shown. Afterpressing OK, the Project Component Map Preview dialog box is produced:

CD-IPE-8 For each Simulator Item (unit or block), the Current Map List shows allcorresponding equipment items in Aspen IPE. Observe that for the default configuration,Standard-Total, five equipment items are included: TW-TRAYED (tower), HE FIXED T-S(condenser), HT HORIZ-DRUM (reflux accumulator), CP CENTRIF (reflux pump), andRB U-TUBE (reboiler). Note that the two C entries denote stream splitters. Note also thatto include a reboiler (bottoms) pump, a distillate pump, and two product heat exchangers,the configuration is switched from Standard-Total to Full-Single. For this example, areboiler pump will be added, as discussed in the section on Adding Equipment.Furthermore, each equipment item has a specific type assigned by Aspen IPE that can bemodified. To modify the equipment type, highlight the item to be modified. In thisexample, the kettle reboiler with U-tubes is replaced by a kettle reboiler with a floatinghead. To begin, the RB U-TUBE reboiler is deleted by highlighting it and pressing DeleteOne Mapping:

CD-IPE-9New Mapping is pressed and reb is highlighted on the screen that appears. Then, OK ispressed.

Reboiler is chosen from the dialog box that appears, and finally a Kettle type reboilerwith floating head is selected as the last step of the replacement procedure.

CD-IPE-10After these steps are completed, the modified mapping should appear on the ProjectComponent Map Preview dialog box:

Other mappings can be altered in a similar fashion. For example, for the condenser, themapping is altered from a shell-and-tube heat exchanger with a fixed tube sheet to onewith a floating head. When the desired changes are completed, press OK to continue andwait for the equipment mapping and sizing to be completed.

CD-IPE-11 At this point, the equipment items have been sized by Aspen IPE (because theSize ICARUS Project Components button was checked in the Map dialog box), whosecalculations are based upon the simulator data, as well as the default values specifiedearlier. As each equipment item is sized, it appears in the Aspen IPE Main window as alist; that is, the List window. Note that the Project Explorer window displays the ProcessView:

The blue boxes to the left of each item in the list indicate the Project Components.The yellow arrows inside the boxes indicate that the equipment item was obtained fromthe mapping of a process simulation unit, whose name appears after its box. Note that bydefault Aspen IPE lists all of the equipment items in the Workbook Mode, as shownabove. The List tab at the bottom of the Main window denotes that the equipment itemsare listed in the Workbook Mode. Also note that user-inputted equipment items, such as areboiler pump (not included in the above frame), are represented in the Workbook by blueboxes without the yellow arrow. To add these equipment items, see the section AddingEquipment. The OK in the Status column of the Workbook indicates that the minimumrequired information for costing the equipment is available. When one or more items aremissing, a question mark appears instead, alerting the user to provide a specification(s) sothat the equipment-sizing step can proceed.

In addition, it is possible to view the IPE Process Flow Diagram. This is

accomplished using the View pulldown menu and clicking on Process Flow Diagram toproduce:

CD-IPE-12Note that the unit icons and streams have been repositioned using “drag and drop”facilities. It is also possible to view a list of the process streams utilized by Aspen IPE;that is, a list of all streams and their physical properties in the Process Flow Diagram.Using the View pulldown menu, click on Streams List to produce:

CD-IPE-13Finally, the IPE Block Flow Diagram shows the simulation flowsheet. It is displayedusing the View pulldown menu and clicking on Block Flow Diagram to give:

Mapping Results. After Aspen IPE has mapped and sized the equipment items, itis prudent to check the results, especially for major equipment items such as towers,compressors, and chemical reactors. These items are usually very expensive, andconsequently, it is a good practice to estimate equipment sizes independently forcomparison with the Aspen IPE results. To view the Aspen IPE results for an equipmentitem, double click on the item on the IPE Workbook window or on its icon in the ProcessFlow Diagram. For example, the following component specification form, whichcontains some of the sizing results, is obtained for the depropanizer tower.

CD-IPE-14 Observe that the depropanizer tower was designed by Aspen IPE to have a 5 ftdiameter and a 42 ft (tangent-to-tangent) height using sieve trays. Note that the numberof trays is the number of equilibrium stages (12 = 14 – 2, excluding the condenser andreboiler) divided by the tray efficiency (0.8), which is 12/0.8 = 15. With a 2-ft trayspacing, a 4-ft high disengagement region at the top and a 10-ft high sump at the bottom,the nominal vessel tangent-to-tangent height is 2 × 14 + 4 + 10 = 42 ft, as shown byAspen IPE. Also, Aspen IPE calculated a design temperature and pressure in accordancewith the Design Criteria specifications, used the default shell material, A515 (which iscarbon steel for pressure vessels at intermediate and higher temperatures), and used thedefault tray material, A285C (which is for carbon steel plates in pressure vessels thathave low and intermediate strength). Material codes, alloy types, and maximum servicetemperatures are tabulated in the chapter on Materials Selection in the ICARUS ReferenceManual (press the Help button in Aspen IPE and follow the path Aspen Icarus ProcessEvaluator 11.1 → Icarus Reference). Furthermore, the effect of material on size and costcan be determined easily. In some cases, a high-strength alloy, that is more expensive perpound, may have thinner walls and be less expensive than a low-strength material that isless expensive per pound.

Changes can be made to any of the equipment sizes computed by Aspen IPE or tothe default values used by Aspen IPE. As changes are made, dependent results areadjusted by Aspen IPE.

A more detailed report can be obtained in two ways. First, right click on theequipment item in the Process Flow Diagram and select Item Report in the menu thatappears. Alternatively, right click on the equipment item in the Project View of the

CD-IPE-15Project Explorer in the Main window (or in the List View) and select Item Report in themenu that appears. These steps produce the Item Report, a portion of which is illustratedhere for the condenser:

Note that only a small portion of the Item Report is shown above. The rawsurface area, 9,652 ft2, is quite large because the log-mean temperature difference,12.7°F, is relatively small. This is related to the condenser pressure which was set at 248psia. At this pressure, the distillate enters the condenser at 125°F and leaves as asaturated vapor at 115°F. Using cooling water heated from 90 to 120°F, the small log-mean temperature difference is obtained. It might be preferable to increase the columnpressure to increase the log-mean temperature difference and reduce the condenser area.However, at a higher pressure the separation would become somewhat more difficult,resulting in more trays. Note that Aspen IPE can easily compare the capital costs atvarious pressures. Note also that Aspen IPE used two floating-head, shell-and-tube heatexchanger in parallel for condensing the overhead vapor. Each condenser has two tubepasses with a temperature correction factor [FT in Eq. (13.7)] of 0.635. The number oftube and shell passes for each exchanger can be seen on the report produced by doubleclicking on the condenser in the IPE Workbook window or on the condenser icon in theProcess Flow Diagram. It might be possible to improve the condenser design by re-sizing the unit with different numbers of shell and tube passes to give a correction factorclose to unity.

CD-IPE-16Standard Basis

Six standard basis profiles are available within Aspen IPE for estimating thecapital cost. These model the nature of the contractor to execute the project, dependingon the size of the project, as shown below. Three of the profiles are for projects to beexecuted by an Owner company (0, 1, and 2), and the other three are for projects to beexecuted by Engineering and Construction firms (3, 4, and 5). For the smalldepropanizer project of this example, the LOCAL CONTRACTOR is appropriate.

To select a standard basis profile for a project, in the Project Basis view, right-click onthe Basis for Capital Costs. Click Select to choose the most appropriate profile.

The Basis for Capital Costs includes specifications for process controls, plantlocation, currency, wage rates, units of measure, and contractor profiles. Default valuesare provided for all entries, most of which need not be adjusted.

When modifying the Basis for Capital Costs, changes can be made to the GeneralStandard Basis Specifications or to the Construction Workforce and Indexing. To viewthe General Standard Basis Specifications, the Project Basis tab is selected in the ProjectExplorer. Double-click the General Specs entry under the Basis for Capital Costsheading to produce the Standard Basis-IP dialog box:

CD-IPE-17 For the depropanizer column, most of the default values are acceptable. Becausea single distillation system would be installed normally on an existing plant site, usingutilities provided by the site, the Project Type would not be selected as GrassRoots/Clear field. This Project Type would cause new items, already provided at the site,to be included in the design and cost estimates. Typically, these include a new controlsystem and electrical substation components.

Under Project Type, click on the Value field to produce a pulldown menu thatdisplays the options:

Grass Roots/Clear field

While guidelines are not provided concerning the selection of Project Type, costs can becomputed for each option, if desired. Through examination of the results, the defaultvalues and items included or omitted can be observed. When selecting Plant addition –suppressed infrastructure, items involving the new control system, electrical switchgear,and transformers, are not provided. These are not needed for the addition of thedepropanizer column to an existing process.

CD-IPE-18Equipment Costing

Aspen IPE estimates the purchase and installed cost of each equipment itemindividually or provides estimates for all of the equipment items (i.e., the entire project)using a single command. For an individual unit, right click on the unit in the List Viewand select Evaluate Item. Aspen IPE produces a detailed item report for the unit. For thedepropanizer tower, by scrolling about a third of the way down the report, the followingsummary of the cost estimates appears:

Observe that the tower designed by Aspen IPE has a Purchased (Equipment and Setting)Cost of $64,100 and an Installed Direct Cost of $192,600, which includes the cost of thetower and setting it in place on its foundation (civil). At this point, the designer canobserve the effects of modifications in the design specifications on these costs for theunit. Be aware that the Total Material and Manpower Cost is the cost of the equipmentitem and the direct cost of installation materials and labor (directly related to theequipment item). These include the piping and field instruments that bring the processstreams to and from the tower; the foundation to support the tower, structural steel (e.g.,ladders and platforms attached to the tower); electrical lighting, heat tracing, cable, andlocal components; insulation; piping; and fireproofing. It does not include: (1) thefractional cost of buildings, pipe racks, the project control system or electricalsubstations, fire control systems, chemical and storm sewers and drains, treatmentsystems, fences, guard houses, etc.; (2) the work required to perform basic and detailengineering, to procure all project components, and to manage the engineering process;and (3) taxes, freight to the site, permits, royalties, etc.

CD-IPE-19 Consequently, the total material and manpower cost is not the total bare modulecost discussed in Section 16.3 of the textbook. The estimate reported by Aspen IPE doesnot include contractor engineering costs, indirect costs, cost of pipe racks and intra-plantpiping, and the cost of sumps and sewers, which can be added to the project as additionalitems. Furthermore, because the report focuses on an equipment item and its associatedinstallation items and costs, materials and manpower items not typically charged to thetower (e.g., charges for instrument testing, pipe testing, and equipment grounding) areexcluded. These costs are accumulated for each area that contains project componentsand are summed for the entire project, as discussed later in this section.

To have Aspen IPE estimate the capital costs of all the units at once (i.e., theentire project), press the Evaluate Project button on the IPE Main window. The EvaluateProject dialog box appears. The dialog box shows the default report file name,CAP_REP.CCP. The contents of this report are viewed in the ICARUS Editor. If youprefer a different name, e.g., DEC3 as shown below, enter it in the Report File field.

When finished with its evaluation, Aspen IPE displays a window that contains anexecutive summary of its results. This window is not shown here. Note that when theuser presses the Tools pulldown menu, selects Options, and then View Spreadsheet inExcel, Aspen IPE is activated to prepare several spreadsheets, including the EquipmentSummary, Utility Summary (available in Version 12.1), ProjSum, Executive Summary,and Run Summary spreadsheets. For the details of these spreadsheets, see the AspenIPE User’s Guide (press the Help button and follow the path Aspen Icarus ProcessEvaluator User’s Manual → Evaluating the Project → Reviewing Investment Analysis).

To view a detailed report of the capital costs, access the ICARUS Editor bypressing the Capital Costs ($) button on the IPE Main window. On the Select ReportType to View dialog box, mark the Evaluation Reports checkbox and press the OK button.Note that when the Interactive Reports checkbox is pressed, the Aspen ICARUS Reporterdialog box is produced. This permits the user to select individual items to be examinedrather than entire reports as discussed below.

CD-IPE-20 The ICARUS Editor displays the report in two adjacent windows, with the majorsubject headings listed in the left-hand window. Most of this information, thoughnecessary for obtaining accurate cost estimates, is far too detailed for most estimatesduring the conceptual design stage, and hence, is normally not printed by processengineers, for whom these course notes are intended. Of greatest interest to processengineers, is the information in the following two sections:

1. Equipment List 2. CONTRACTOR NO. 1 PRIME CONTRACTOR

which are accessed by double-clicking on these titles in the left-hand window. It is

recommended that just small portions of the report be printed. This is accomplished byhighlighting the desired section and pressing the Print button on the toolbar. It is oftenpreferable to print in landscape format.

When the appropriate specifications are made, Aspen IPE computes annualoperating costs, as well as a complete profitability analysis, the results of which appear inthis Investment Analysis spreadsheet. These notes discuss capital cost estimation onlybecause the spreadsheet, Profitability Analysis-1.0.xls, which is discussed in Section 17.8of the textbook, is used to compute operating costs, working capital, and profitabilitymeasures.

As shown below, the List of Equipment and Bulk Material by Area portion of thereport is displayed when the Equipment List is accessed. This provides the Purchased(Equipment & Setting) and Installed Direct Costs (i.e., Total Material and ManpowerCost or Total Direct Materials and Labor Cost) for each piece of equipment, e.g., thereboiler as shown next. Note that the right-hand window below is displayed using a 7-point font. This is achieved by pressing the Select Font button on the toolbar.Furthermore, portions of the complete printed output are provided in Appendix III ofthese notes.

CD-IPE-21 In summary, the equipment sizes, purchase costs, and total material andmanpower cost for the depropanizer system (without the reboiler pump) are as follows:

D1 Tower 5.0 ft diam. 64,100 192,600

TOTAL $280,300 $646,800

The Contract Summary section of the Capital Estimate Report is displayed whenthe CONTRACT NO. 1 PRIME CONTRACTOR is accessed. The entries shown below arein a 6-point font and are totals for all of the equipment items (i.e., the entire project).Note that selected portions of the complete printed output are provided in Appendix III ofthese notes.

CD-IPE-22 Note that the entry for the purchased equipment, $289,200, from line 1, isapproximately the sum of the entries for the pieces of equipment provided above,$280,300. The difference is due to the Misc. Item Allowance ($8,500) and the WarehouseSpares ($370). These additional items are in Code of Accounts 105 and 107 and appearin the Code of Accounts Summary section of the Capital Estimate Report (just below theContract Summary.)

The total direct material and manpower costs for construction of the plant are$605,400 and $152,100, as shown in line 11. These sum to $757,500 and include itemsthat cannot be charged to the individual equipment items (e.g., charges for instrumenttesting, pipe testing, and equipment grounding). Note that the installed costs of theequipment items are displayed on the List View:

CD-IPE-23The installed costs sum to $646,800; that is, $108,700 less than the total direct cost ofmaterials and manpower for installation of the plant, $757,500. This Installed DirectCost, CDI, is referred to in Chapter 16 of the textbook as the Total Direct Materials andLabor Cost, CDML. Finally, the materials and manpower items that are not chargeable tothe individual equipment items are displayed in the Area Bulk Report within the CapitalEstimate Report:

AREA PIPE TESTING 0. 192 4564. 4564.

GRADE UNPAVED AREA 7534. 127 2648. 10182.

AREA INSTRUMENT TESTING 0. 95 2124. 2124.

AREA INSTR. RUNS,TRAYS,JBOX. 3086. 60 1266. 4352.

AREA EQUIPMENT GROUNDING 185. 11 231. 416.

AREA PILED FOUNDATION 8807. 83 1407. 10214.

Number of piles 14

AREA ELECTRICAL TESTING 0. 16 344. 344.

AREA ROTATING EQP SPARE PARTS 370. 0 0. 370.

CD-IPE-24These non-chargeable items add to $35,545. Together with the Other item on line 10 ofthe Contract Summary, $56,300, and Code of Accounts item 105, for equipmentcontingencies to allow for design changes, $8,500, these sum to approximately $100,300(which is sufficiently close to $108,700, the difference reported above).

Returning to the Capital Estimate Report, material and manpower costs

associated with G and A (General and Administrative) Overheads, $18,200 and $4,600,are obtained from line 13, and material and manpower charges associated with ContractFees, $21,900 and $16,500, from line 14. These sum to $61,200. The contractorengineering and indirect costs are in row 15, BASE TOTAL, in the first column, underDESIGN ENG’G AND PROCUREMENT K-USD, and in the fifth column, underCONSTRUCTION INDIRECTS K-USD. These are:

Contractor Engineering Costs $383,700

Indirect Costs $365,700

Together with the fees for materials and manpower G and A Overheads and ContractFees, these are added to the total direct installed equipment costs, CDI, to give the IBLTotal Bare Module Cost, CTBM.

Finally, all of the Aspen IPE results can be reproduced using the DEC3 folder (onthe CD-ROM in the Aspen Eng. Suite folder) from within Aspen IPE.

Total Permanent Investment

The total permanent investment is computed by the spreadsheet, Profitability

Analysis-1.0.xls, discussed in Section 17.8 of the textbook. When using the Aspen IPEoption, the user enters:

Total Direct Materials and Labor Costs $757,500

Thus far, all of the equipment items have originated with the simulation unitsfrom an ASPEN PLUS simulation. After the mappings have been completed, yellowarrows are placed in the blue boxes associated with each equipment item in the AspenIPE Main window. Also, in the Process Flow Diagram, all of the streams are yellow,with the exception of the IPE-generated utility streams, which are green. When it isdesirable to add a piece of equipment that is not in a simulation or has not been createdduring the mapping of simulation units by Aspen IPE, the following steps are taken.From the IPE Main window, press the Project View tab at the bottom of the left-handwindow (i.e., the Project Explorer window) to give:

CD-IPE-25Then, highlight Main Project, right click, and press Add Area to produce the AreaInformation dialog box in which an Area Name (e.g., New Item) is entered with itsdimensions. Here, a 50’x 50’ area is reserved and used to estimate piping lengths, etc.This is adequate for most applications. Note that the original area for the plant, whichwas named Miscellaneous Flowsheet Area by Aspen IPE, is also 50’ x 50’ by default.

Press OK and the new area, which is named New Item, appears on the Project View (leftwindow) of the IPE Main window.

CD-IPE-26Next, highlight the New Item area, right click, and click on Add Project Component toproduce the ICARUS Project Component Selection dialog box. For the addition of areboiler pump, enter Reboiler Pump as the Project Component Name, highlight Processequipment and press the OK button.

Continue through the appropriate menus until the desired equipment type is obtained,which in this example is a centrifugal pump.

CD-IPE-27CD-IPE-28After the OK button is pressed, the pump specification form is displayed.

CD-IPE-29Note that the specifications are incomplete because the Reboiler pump has not beenconnected into the main process, which resides in the Miscellaneous Flowsheet Area, asshown in the IPE Process Flow Diagram:

The Reboiler pump is positioned in the upper-left-hand corner of the Process FlowDiagram in the New Item area, independent of the Miscellaneous Flowsheet Area.Observe that the Reboiler pump appears in the New Item area on the Project View.

Before proceeding, after completing this example, it was brought to our attentionthat reboiler pumps are used normally with vertical reboilers, not with kettle reboilers.When appropriate to add a reboiler pump, or any other equipment item, to the mapping,the procedures in this section should be followed.

To insert the Reboiler pump into the liquid stream from the sump, ICP-BE, pressthe Edit Connectivity button and place the cursor over the Reboiler pump, after which thecursor becomes a hand. Keeping the left-mouse button depressed, drag the Reboilerpump over the ICP-BE stream. Release the mouse and click with the left-mouse button toinsert the Reboiler pump. After the streams are realigned, the Process Flow Diagramappears as follows:

CD-IPE-30Note that a new stream, which appears in white, has been created and named ICP-BE_2by Aspen IPE.

Although the Reboiler pump has been inserted into the process, it remains in theNew Item area. To move it into the Miscellaneous Flowsheet area, in the Project View,drag and drop the Reboiler pump from the New Item Area to the Miscellaneous Flowsheetarea. This results in:

CD-IPE-31 Next, right click on the Reboiler pump and select Size Item on the menu thatappears. After the pump is sized, double click on the pump icon to display thecomponent specification form:

Note that the design capacity of the Reboiler pump has been adjusted to 765.5 gpm,which is 10 percent higher than the flow rate leaving the sump, a default specification inthe Design Criteria. At this point, a fluid head of 20 ft is entered, which should besufficient to convey the bottoms liquid to the reboiler. To obtain the variables for theICP-BE stream, double-click on it:

CD-IPE-32Observe that 765.5 gpm is 10 percent higher than 695.9 gpm, which is equivalent to theliquid mass flow rate, 160,645 lb/hr. When the effluent stream, ICP-BE_2, is clicked on,the stream report does not display the stream properties because the stream has beenreferenced to the ICP-BE stream.

This procedure is repeated to add other equipment items, which may be added tothe New Item area or to other new areas.

To estimate the installed cost of the Reboiler pump, either right click on Reboilerpump in the Project View or on its icon in the Process Flow Diagram. Then, selectEvaluate Item. A brief report that contains the installed cost, $44,700, can be accessed byhighlighting Reboiler pump in the Project View and pressing the List tab to obtain theWorkbook. A complete report is obtained by re-evaluating the capital estimates for theprocess. This is accomplished by pressing the Evaluate Project button and requestingthat all equipment items be re-evaluated. The detailed report appears in the CapitalEstimate Report in the List of Equipment and Bulk Material by Area section. It can beaccessed by selecting Equipment List under Miscellaneous Flowsheet in the left-handwindow:

Note that no equipment items remain in the New Item section of the report.

Having added the Reboiler pump, the total permament investment can be re-estimated as discussed in the prior section. This discussion is not repeated here.

Finally, all of the Aspen IPE results can be reproduced using the DEC3RP folder(on the CD-ROM in the Aspen Eng. Suite folder) from within Aspen IPE.

CD-IPE-33Applying Alternative Utilities

When desired, the default utility applied by Aspen IPE can be altered interactivelyfor a particular equipment item, such as a condenser or reboiler, after it has been mapped.For example, when the resulting surface area of a reboiler is too large due to a small log-mean-temperature-difference, the steam utility can be replaced with steam at a higherpressure to reduce the area, being careful to stay in the nucleate boiling region.

This is illustrated for the reboiler of the depropanizer as an example. For thisreboiler, Aspen IPE uses steam at 50 psi as the default utility. To change to higher-pressure steam, say at 100 psi, the following steps are taken.

In the Process View or Process Flow Diagram, right click on the reboiler andselect Re-Size Item from the menu that appears. This produces the Interactive Sizingdialog box, as shown below:

In the Item 1 column that contains the values, the items for Hot Inlet Stream andHot Outlet Stream are ICUST-IN and ICUST-EX, respectively, which correspond to thedefault utility, in this case, steam at 50 psi. To change to steam at 100 psi, right click onthe appropriate cells and select Steam @ 100 PSI – IPE Utility from the pull-down menuthat appears. Next, delete the Final Surface Area, previously computed, since it must bere-sized by Aspen IPE:

CD-IPE-34When OK is pressed, the reboiler is re-sized.

After the reboiler is re-sized, right click on the reboiler again and select ItemReport from the pop-up menu. In the Sizing Data section, the new results for the reboilerare displayed:

CD-IPE-35Using steam at 100 psi, the surface area is 1,262 ft2, reduced from 3,580 ft2, while thelog-mean-temperature-difference is 72.7°F, increased from 25.7°F.

Finally, the capital cost of the entire process is re-evaluated since the cost of thesmaller reboiler is lower. This is accomplished by pressing the Evaluate Project buttonon the toolbar and selecting Evaluate All Items. The results appear in the CapitalEstimate Report in the List of Equipment and Bulk Material by Area section. They areaccessed by selecting Equipment List under Miscellaneous Flowsheet in the left-handwindow:

These steps are repeated when it is desired to change the default utilities for otherequipment items in the process.

Furthermore, for most equipment items, other specifications can be adjusted usinginteractive sizing. This can be accomplished for condensers, reboilers, flash drums,reflux accumulators, storage vessels, pumps, and compressors. Note, however, thatinteractive sizing is not possible for reactor vessels. For a complete listing of equipmentitems that can be sized interactively, refer to the chapter on Sizing Project Components inthe Aspen IPE User’s Guide (Aspen Icarus Process Evaluator User’s Manual → SizingProject Components).

CD-IPE-36MONOCHLOROBENZENE SEPARATION PROCESS

In this section, equipment sizes and costs are estimated for themonochlorobenzene (MCB) separation process, which is discussed in Section 4.4 of thetextbook and in the multimedia portion of the CD-ROM (ASPEN → Principles ofFlowsheet Simulation → Interpretation of Input and Output → Sample Problem) thatcontains these course notes. Beginning with the file, MCB.bkp, which is available on theCD-ROM, additional mixture properties are added and the DISTL subroutine, used tomodel the D1 distillation column, is replaced with the RADFRAC subroutine. The refluxratio computed using the RADFRAC subroutine is 3.35, as compared with 4.29computed using the approximate DISTL subroutine. Also, the stream flow rates differslightly (< 1%). Both of the files, MCB-IPE.bkp and MCB-IPE.rep, are on the CD-ROM.

Initial Setup

After sending the file, MCB-IPE.rep to Aspen Icarus, the user is ready to useAspen IPE. Aspen IPE is opened automatically and the Create New Project dialog boxappears. After the Project Name MCB is entered, the Inch-Pound (IP) unit set is selectedin the Project Properties dialog box.

After OK is pressed, Aspen IPE loads the information associated with eachprocess model in ASPEN PLUS. When completed, the IPE Main window appears:

CD-IPE-37 The MCB separation process has two types of columns, an absorber and adistillation column, each having a distinct tray efficiency. Absorber efficiencies arenormally low, at roughly 20%, while efficiencies for distillation columns areconsiderably higher, in this case at about 60%. This difference must be taken intoaccount when proceeding with Aspen IPE.

Because Aspen IPE allows only one specification for the tray efficiency, it isnecessary to map and size each of the columns separately, with the appropriate efficiencyspecified in the Design Criteria prior to each mapping. Note that in the Design Criteria-IP dialog box, the parameters for trayed towers, including the tray efficiency are near thebottom of the list:

To size the absorber column (A1-block), a tray efficiency of 0.2 (or 20%) is entered. Noother changes to the default values are necessary.

Changes to the Utility Specifications, such as the cooling water temperatures, aremade at this point.

CD-IPE-38Mapping Process Simulation Units to Aspen IPE

To map a single process unit, right-click on the selected item on the Aspen IPEMain window, and choose Map. In the Map dialog box, select Map Selected Item(s), anduse Default and Simulator Data as the basis:

Press OK to produce the Project Component Map Preview dialog box (not shown here).Since the Current Map List does not need to be altered, select OK to map the A1 unit.When the mapping and equipment sizing has been completed, the A1 unit has beenadded to the list of Project Components, as shown below:

CD-IPE-39 Before mapping the distillation unit, D1, the tray efficiency is changed to 0.6 inthe Design Criteria. Subsequently, each of the remaining equipment items is mappedand sized, one at a time, as described above. Note that the unit H1 is too small to bemapped as a floating-head heat exchanger. Consequently, it is necessary to change thedefault equipment type to a Double-pipe heat exchanger, which is more appropriate forthis application. To change the mapping, select HE FLOAT-HEAD in the Current MapList and press the Delete One Mapping button:

After these steps are completed, the Current Map List is modified in the ProjectComponent Map Preview dialog box:

Note that when the sizing calculations are being carried out for the flash vessel,F1, two Message dialog boxes appear. The first indicates that the diameter is calculatedto be 2.007 ft, but that the user-specified minimum value of 3 ft is used instead. Thesecond indicates that the L/D ratio is 1.67, rather than 3.0 from the Design Criteria.

CD-IPE-41Also, for the heat exchanger, H1, a 1-degree difference between the inlet and outlettemperatures of the hot stream is assumed. The unit M1 is a mixing junction betweentwo pipes and the unit S1 is a simple pipeline splitter. Size and cost estimates are notneeded for these units. The unit T1 represents a treater, which is not being considered atthis point in the design of the MCB separation process.

Aspen IPE maps the mixer M1 and splitter S1 as Quoted Items with zero cost.The default mapping for the treater T1 is a VT CYLINDER, with size and cost estimatescomputed. This default mapping is replaced with a Quoted Item having zero cost. Toaccomplish this, delete the mapping for T1. In the Project View, right click on T1, thenon Map. On the Map dialog box, click on OK to produce the Project Component MapPreview dialog box. Delete the VT CYLINDER mapping and click on New Mapping, toproduce the ICARUS Project Component Selection dialog box. Click on ProjectComponents, select Quoted equipment, and click OK. This places the unit T1 into theList View with a C, to indicate that it is a Quoted Item having zero cost.

After all of the equipment items have been mapped and sized successfully, theIPE Main window is displayed:

CD-IPE-42 When the mapping and sizing are completed it is prudent to check the equipmentsizes computed by Aspen IPE, especially for major equipment items such as towers, largeheat exchangers, compressors, and chemical reactors. For the MCB separation process,the two towers are of particular interest. To view the Aspen IPE result for an equipmentitem, double click on the item of interest in the IPE Main window. For the absorber, thisproduces the following results:

CD-IPE-43Note that the column is designed to have a 1.5 ft diameter, a 42 ft (tangent-to-tangent)height, and 15 trays, in accordance with the specifications in Figure 4.23 of the textbook(Seider et al., 2004). Because of the small diameter, a packed column would bepreferred, but is not considered here.

CD-IPE-44Similarly, the distillation column is designed to have a 3 ft diameter, a 72 ft (tangent-totangent) height, and 30 trays, also in accordance with Figure 4.23 in the textbook (Seideret al., 2004).

Standard Basis

As for the depropanizer discussed earlier, the MCB separation process can beviewed as representing an addition to an existing plant. Consequently, the standard basisprofile is selected to be LOCAL CONTRACTOR and the Project Type is selected as Plantaddition – suppressed infrastructure.

Equipment Costing

Aspen IPE estimates purchase and installed costs for the equipment unitsindividually or for the entire project using a single command. For the MCB separationprocess, it is convenient to have Aspen IPE estimate the costs for the entire project atonce. After pressing the Evaluate Project button on the IPE Main window, the EvaluateProject dialog box appears:

As discussed for the depropanizer, Aspen IPE prepares the Capital Estimate Report,MCB.ccp, which contains detailed listings of the items to be procured to install theequipment (classified in the areas of piping, instrumentation, electrical, structural steel,and insulation), estimates of the man-hours required for installation, estimates of thecosts, and an installation schedule. Estimates for contractor engineering and indirectcosts are listed as well.

The ICARUS Editor displays the report in two adjacent windows, with a listing ofthe major subject headings listed in the left-hand window. Most of this information,though necessary for obtaining accurate cost estimates, is far too detailed for mostestimates made in the conceptual design stage, and hence, is normally not printed byprocess engineers, for whom these notes are intended. Of greatest interest to processengineers, is the information in the following two sections:

CD-IPE-45 1. Equipment List 2. CONTRACTOR NO. 1 PRIME CONTRACTOR

which are accessed by double-clicking on these titles in the left-hand window. It is

recommended that just small portions of the report be printed. This is accomplished byhighlighting the desired section and pressing the Print button on the toolbar. It is oftenpreferable to print in landscape format.

As shown below, for the absorber, the List of Equipment and Bulk Material byArea portion of the report is displayed when the Equipment List is accessed. Thisprovides the Purchased (Equipment & Setting) and Installed Direct Costs (i.e., TotalMaterial and Manpower Cost or Total Direct Materials and Labor Cost) for each pieceof equipment. Note that the right-hand window below is displayed using a 7-point font.This is achieved by pressing the Select Font button on the toolbar. Furthermore, portionsof the complete printed output are provided in Appendix IV of these notes.

In summary, the equipment sizes, purchase costs, and total material andmanpower cost for the MCB separation process are tabulated below:

A1 Tower 1.5 ft diam. 16,000 110,000

D1 Tower 3.0 ft diam. 53,500 179,200

H1 Heat exchanger 161 ft2 16,100 58,100

H2 Heat exchanger 196 ft2 12,400 52,900

F1 Flash vessel 264 gal 7,100 54,200

TOTAL $154,400 $671,900

The Contract Summary section of the Capital Estimate Report is displayed whenthe CONTRACT NO. 1 PRIME CONTRACTOR is accessed. The entries shown below arein a 6-point font and are totals for all of the equipment items in the project. Note thatportions of the complete printed output are provided in Appendix IV of these notes.

CD-IPE-47 Note that the entry for the purchased equipment, $159,500, from line 1, isapproximately the sum of the entries for the pieces of equipment provided above,$154,400. The difference is due to the Misc. Item Allowance ($4,700) and the WarehouseSpares ($430). These additional items are in Code of Accounts 105 and 107 and appearin the Code of Accounts Summary section of the Capital Estimate Report (just below theContract Summary.)

The total direct material and manpower costs for construction of the plant are$536,200 and $249,500, as shown in row 11. These sum to $785,700 and include itemsthat cannot be charged to the individual equipment items (e.g., charges for instrumenttesting, pipe testing, and equipment grounding). Note that the installed costs of theequipment items are displayed on the List View:

CD-IPE-48These installed costs sum to $671,900; that is, $113,800 less than the total direct cost ofmaterials and manpower for installation of the plant, $785,700. This Installed DirectCost, CDI, is referred to in Chapter 16 of the textbook as the Total Direct Materials andLabor Cost, CDML. Finally, the materials and manpower items that cannot be charged tothe individual equipment items are displayed in the Area Bulk Report within the CapitalEstimate Report: A R E A B U L K R E P O R T

AREA PIPE TESTING 0. 328 7771. 7771.

GRADE UNPAVED AREA 7534. 127 2648. 10182.

AREA INSTRUMENT TESTING 0. 179 4012. 4012.

AREA INSTR. RUNS,TRAYS,JBOX. 3471. 68 1450. 4921.

AREA EQUIPMENT GROUNDING 369. 23 462. 831.

AREA PILED FOUNDATION 13840. 131 2211. 16051.

Number of piles 22

AREA ELECTRICAL TESTING 0. 20 430. 430.

AREA ROTATING EQP SPARE PARTS 430. 0 0. 430.

CD-IPE-49These additional costs sum to approximately $48,902. Together with the Other item online 10 of the Contract Summary, $49,900, and Code of Accounts item 105, forequipment contingencies to allow for design changes, $4,700, these sum to approximately$103,500 (which, for profitability analysis in the conceptual design stage, is sufficientlyclose to $113,800, the difference reported above).

Returning to the Contract Summary, material and manpower costs associated withG and A (General and Administrative) Overheads, $16,100 and $7,500, are obtainedfrom line 13, and material and manpower charges associated with Contract Fees, $20,400and $25,700, from line 14. These sum to $69,700. The contractor engineering andindirect costs are in row 15, BASE TOTAL, in the first column, under DESIGN ENG’GAND PROCUREMENT K-USD, and in the fifth COLUMN, under CONSTRUCTIONINDIRECTS K-USD. These are:

Contractor Engineering Costs $558,300

Indirect Costs $482,600

Together with the fees for materials and manpower G and A Overheads and ContractFees, these are added to the total direct installed equipment costs, CDI, to give the IBLTotal Bare Module Cost, CTBM.

Finally, all of the Aspen IPE results can be reproduced using the MCB folder (onthe CD-ROM in the Aspen Eng. Suite folder) from within Aspen IPE.

Total Permanent Investment

The total permanent investment is computed by the spreadsheet, Profitability

Analysis-1.0.xls, discussed in Section 17.8 of the textbook. When using the Aspen IPEoption, the user enters:

Total Direct Materials and Labor Costs $785,700

When a new project is created within Aspen IPE, a folder having the project name(e.g., DEC3) is created in the Program Files|Aspen Tech\ Aspen Icarus 11.1\Data\Archive_IPE folder. As work with Aspen IPE proceeds, various files are createdand stored in this project folder; for example, the DEC3.ccp file, which contains theCapital Estimate Report for the depropanizer.

When returning to work with Aspen IPE, using the File pulldown menu, open thefolder having the appropriate project name. This produces the Open an IPE Projectdialog box. Select the Project Name and press the OK button. This produces theProcess/Project View window; that is, the IPE Main window.

When working in the Process/Project View window, to examine any portion of

the Capital Estimate Report (which is automatically stored in your Projects folder after ithas been generated), press the Capital costs button ($) on the toolbar. This produces theSelect Report Type To View dialog box. Select one of the two options to have Aspen IPEdisplay the capital cost report as an HTML file or in the ICARUS Editor. Then, pressOK. Note that when more than one report file exists, the Select Capital Cost Report Filedialog box is produced, from which the appropriate report file is selected. This producesthe Capital Estimate Report.

It is also possible to examine a .ccp file using the NETSCAPE or EXPLORER

browser by double-clicking on the appropriate file, which has a browser icon, in itsassociated project folder. Each item in the contents that is produced provides a link to itssection of the Capital Estimate Report.

**** PROFILES ****

**NOTE** REPORTED VALUES FOR STAGE LIQUID AND VAPOR RATES ARE THE FLOWS FROM THE STAGE EXCLUDING ANY SIDE PRODUCT. FOR THE FIRST STAGE, THE REPORTED VAPOR FLOW IS THE VAPOR DISTILLATE FLOW. FOR THE LAST STAGE, THE REPORTED LIQUID FLOW IS THE LIQUID BOTTOMS FLOW.

Environmental – Soil Treatment

Phytoremediation of Lead-Contaminated Sites A-II.9.1

Soil Remediation and Reclamation A-II.9.2

Fuel Processor for 5 KW PEM Fuel Cell Unit A-II.10.1

Combined Cycle Power Generation A-II.10.2

Production of Low-Sulfur Diesel Fuel A-II.10.3

Waste Fuel Upgrading to Acetone and Isopropanol A-II.10.4

Conversion of Cheese Whey (Solid Waste) to Lactic Acid A-II.10.5

Ethanol from Corn Syrup A-II.10.6

CD-A-II-3This appendix contains the problem statements for 50 design projects, each prepared for designteams of three students at the University of Pennsylvania by chemical engineers in the localchemical industry. At Penn, each team selects its design project during the first lecture course inthe fall, and spends the spring semester completing the design. In the spring, each group meetsregularly with its faculty advisor and industrial consultants, including the individual who providedthe problem statement, to report on its progress and gain advice.

The problem statements in the file, Design Problem Statements.pdf, on the CD-ROM are in theiroriginal forms, as they were presented to the student design teams on the date indicated. Someprovide relatively little information, whereas others are fairly detailed concerning the specificproblems that need to be solved to complete the design. The reader should recognize that, in nearlyevery case, as the design team proceeded to assess the primitive problem statement and carry out aliterature search, the specific problems it formulated were somewhat different than stated herein.Still, these problem statements should be useful to students and faculty in several respects. Forstudents, they should help to show the broad spectrum of design problems that chemical engineershave been tackling in recent years. For the faculty, they should provide a basis for similar designprojects to be created for their courses.

In formulating design problem statements, the industrial consultants strive to create processopportunities that lead to designs that are timely, challenging, and offer a reasonable likelihood thatthe final design will be attractive economically. Every effort is made to formulate problems thatcan be tackled by chemical engineering seniors without unduly gross assumptions and for whichgood sources of data exist for the reaction kinetics and thermophysical and transport properties. Inthis respect, this was accomplished in each of the problems included herein; furthermore,successful designs were completed by a student design team for most of these problems.

As seen in the contents, the projects have been assigned to one of the following areas, in somecases arbitrarily: Petrochemicals, Petroleum Products, Gas Manufacture, Foods, Pharmaceuticals,Polymers, and Environmental.

CD-A-II-4Credit is given to each formulator on his problem statement. In addition, the names of thecontributors are listed below with many thanks, as their contributions in preparing these designproblems have been crucial to the success of the design course.

Anticipated Sales (in thousands of pounds)

You currently have a fully depreciated, 1,000-gallon batch reactor that is used to manufactureanother product (Product X). This reactor is made of 316SS, which is sufficiently corrosion-resistant for producing the new product as well. Product X is made in 6,000-pound batches thatrequire 36 reactor hours per batch and is sold at a profit of $0.88 per pound. 100 such batches areproduced annually (not expected to change); the rest of the time the reactor is idle. This reactor isjacketed for heating and uses 175 psig saturated steam. The jacket has a heat-transfer area of 88 ft2and an estimated overall heat-transfer coefficient of 100 Btu/ft2hr°F.

O O

OH OH OH OR + 2 ROH + 2 H2O OH Acid Catalyst OR

O O

Di(3-pentyl) malate is made by batch reaction of malic acid with an excess of 3-pentanol, using 0.1weight percent of an acid catalyst such as sulfuric acid (see reaction above). Water is produced as aco-product and must be removed to drive the reaction to completion. Water and 3-pentanol form alow-boiling azeotrope (see CRC Handbook for data) that forms two liquid phases uponcondensation. A typical process scheme would be to carry out the batch reaction above theazeotrope temperature while condensing the overhead vapors into a decanter, recycling the organiclayer to the reactor and removing the aqueous layer (Figure 1, top). This approach can be used withyour existing reactor. A more sophisticated approach would involve interposing a distillationcolumn between the reactor and the condenser, allowing the alcohol-rich vapors off the reactor tostrip water out of the organic recycle (Figure 1, bottom). When the desired conversion is achieved,the product must be treated with aqueous sodium hydroxide to neutralize the residual acidity (dueboth to the catalyst and the unreacted malic acid). The residual 3-pentanol must be stripped off

CD-A-II-6using vacuum (50 mm Hg) with nitrogen sparge at 120°C. Your R&D group has come up with themass-transfer estimates given in Table 1. Finally, the product must be filtered to remove the saltsof neutralization. Your company currently has no vacuum or filtration equipment.

Table 1. Mass Transfer Data

dx dt ( ) = k L a y * − y where x is the mole fraction of 3-pentanol in the liquid , y* is the vapor phase mole fraction of 3-pentanol in equilibrium with x, and y is the vapor phase mole fraction of 3-pentanol. Assume that the Henry’s law constant for 3- pentanol in the product is 1,200 mm Hg.

Superficial Gas Velocity (scf/ft2,min) 2 5 10 20 50

kLa (1/hr) 0.076 0.12 0.17 0.24 0.37

The required product specifications are:

Residual acidity (prior to neutralization) <0.1N

You are being asked to provide the following:

1. An equipment design for a dedicated batch-reactor system to produce dibutyl malate, including a capital cost estimate for both process options shown in Figure 1.2. A batch ticket for a typical production batch. This will itemize the individual steps the operator will follow to produce the batch, including amounts of materials being added, estimated duration of each step and the safety procedures and precautions that must be followed. It should also specify when samples must be taken and what the criteria are for proceeding to the next step.3. A recommendation to management on whether/when to build the dedicated equipment or use the existing reactor, supported by appropriate financial information.

Key process determinations:

Which process option should you use for a new design – with or without the distillation column? How much heat-transfer surface is required and what heating medium (assume you have saturated steam available at 175 psig for $5 per million Btu)? What type of agitation is needed (horsepower and impeller design)? How long will the reaction take? What is the reaction profile (concentrations and temperature vs. time)? How does the composition of the vapor from the reactor change with time? What ratio of alcohol to malic acid should be charged? What types of process control systems are required to ensure product quality? What are you going to do with the aqueous byproduct and the recovered excess alcohol? Is it worth buying any additional vessels for post-treatment, filtration, storage, etc.? What kind of vacuum system should you purchase?

CD-A-II-7 What equipment will be needed for filtration? What will your overall batch cycle time be?

Costs:

Malic Acid, 1,000 kg supersacks, $2,750 each; 50 lb bags, $78 each

Data & Additional Information:

The viscosity (cP) of the reactor contents can be estimated using the equation 0.00211*exp(2,600/T), where T is in Kelvin. Product density is 1.03 g/cc. Assume that this is also the density of the reactor contents at every point in the reaction. Residual acidity can be measure by titration, requiring 15 minutes to obtain a measurement from the time the sample is taken. Residual alcohol and product purity are measured by chromatography, requiring 45 min from the time the sample is taken.

Use the following reaction rate expressions in your model, treating the two acid groups on eachmalic acid molecule as if they are two separate molecules:

Make the following additional assumptions (and be sure to document additional assumption youmake):

Malic acid completely dissolves in 3-pentanol at 70°C.

The heat capacity of the reactor contents is 0.50 Btu/lb°F throughout the process. Assume that the reaction occurs at atmospheric pressure. Assume that all products of neutralization are insoluble. Assume that during filtration only the resistance of the cake itself is significant. No additional equipment must be purchased to transport or charge the solid malic acid.

CD-A-II-8 VENT

Org

Aqu

AQUEOUS BYPRODUCT

VENT

Org

Aqu

AQUEOUS BYPRODUCT

Figure 1. Reaction Schemes for Di(3-pentyl) Malate Manufacture

Acetaldehyde is a versatile chemical intermediate. It is commercially made via the Wacker

process, the partial oxidation of ethylene. That process is very corrosive, requiring expensivematerials of construction. And like all oxidations, over-oxidation of the ingredient and the productreduce the yield, and convert expensive ethylene into carbon oxides.

Acetic acid, produced from inexpensive methanol, would be a good feedstock, if a selective routeto acetaldehyde could be found. Because of the possible legislation of MTBE out of gasoline, theremay be a worldwide glut of methanol, so any chemicals that use methanol may become much moreeconomically attractive. But the reduction of acetic acid to acetaldehyde is notoriously difficult,because aldehydes are easier than acids to reduce.

However, Eastman Chemical has developed a selective palladium catalyst that gives acetaldehydewith selectivity of up to 86% at 46% conversion. Byproducts formed include ethanol, acetone andethyl acetate, all of which can be sold after purification.

CH 3 − COOH + H 2 → CH 3 − CHO + H 2O (main reaction)

CH 3 − COOH + 2 H 2 → CH 3 − CH 2OH + H 2O

CH 3 − COOH + CH 3 − CH 2OH ↔ CH 3 − COO − CH 2 − CH 3 + H 2O

2 CH 3 − COOH + 2 H 2 → CH 3 − CO − CH 3 + CH 4 + H 2O

Distillation of the product will be complicated by the existence of azeotropes between ethanol andethyl acetate, water and ethanol, and water and ethyl acetate. And the acetic acid-water andacetone-water mixtures are famous for their tangent pinches. Rigorous distillation simulations withthermodynamics that accurately predict each of these azeotropes and pinches will be required tohave confidence in the design.

Your company has asked your group to determine whether this new technology should be used inyour Gulf Coast plant. Your job is to design a process and plant to produce 100 MM lb/yr ofacetaldehyde from acetic acid, which is available on the site. Based on past experience, you knowthat you will have to defend any decisions you have made throughout the design, and the bestdefense is economic justification.

Assume a U.S. Gulf Coast location on the same site as a large chemical plant. Acetaldehyde can besold for $0.48/lb, according to your marketing organization. Acetic acid is available on your sitefor $0.16/lb. However, if MTBE is legislated out of gasoline, that price might drop to $0.12/lb.Test your economics with both prices, and make appropriate recommendations. Hydrogen can bepurchased over the plant fence for $0.50/lb at 200 psig. Ethanol, if 99.95% pure, can be sold (on anexcise tax-free basis) for $2.50/gal; however, the ethanol-water azeotrope can also be sold into the

CD-A-II-10fuel market for $1.60/gal. You may sell either or both grades of ethanol, depending on which ismost economical to produce. Ethyl acetate can be sold for $0.60/lb. Acetone can be sold for$0.20/lb. You will need storage tanks, truck or railcar loading stations, etc., for each byproduct thatyou sell, or you may burn them in the boiler for fuel value. Byproducts sold must also meet normalpurity specs for that chemical. All prices listed are in 2002 dollars.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate.Remember that you will be there for the start-up and will have to live with whatever designdecisions you have made.

References:

U. S. Patent 6,121,498 to Eastman Chemical.

A-II.1.3 Ethylene by Oxidative Dehydrogenation of Ethane

(Bruce Vrana, DuPont, January 2001)

Ethylene is the largest volume organic chemical product, with world production over 50 billionpounds per year. It is normally produced by steam cracking of ethane or heavier hydrocarbons.This process is quite energy and capital intensive.

Dow Chemical has recently applied for a patent on a new process, which may require significantlyless investment. In this process, ethane is passed over a catalyst at very high space velocity(100,000/hr or higher), and reacts with oxygen (exothermically!), producing ethylene in goodselectivity (greater than 80% under some conditions) and high conversion. The selectivity issimilar to that in the conventional steam cracking process, but the conversion is higher. Hydrogenin the feed improves the conversion while minimizing the amount of over-oxidation of thefeedstock.

Because the reaction with oxygen is exothermic, the expensive furnaces of the steam crackingprocess should not be required. Much less coke is produced in this reactor system, according toDow, which should result in a much more operable plant.

Dow has patented both a fixed bed supported catalyst and a fluidized bed reactor. The fluidizedbed has a slightly higher selectivity, and would probably be easier to manage the heat load than theless expensive fixed bed reactor. You should use economics and technical criteria to guide yourdecision about which reactor technology to use in the plant design, and discuss this major decisionin your report.

Your company has 1 MMM pounds per year of ethane, which is currently being produced at yourGulf Coast plant and sold for $0.07/lb in 2000. Your team has been asked to evaluate the economicviability of the Dow process for your plant, as a way of upgrading your product and increasing yoursales revenue. Your job is to determine the economic optimum design, maximizing the net present

CD-A-II-11value (NPV) of the project. You may consume all or part of the ethane, which is available. Basedon past experience, you know that you will have to be able to defend any decisions you have madethroughout the design, and the best defense is economic justification. Your plant design must bebacked up with a rigorous simulation of the entire process, with all recycle loops closed.

Your marketing organization believes they can sell ethylene for $0.25/lb in 2001 dollars. Pipelineoxygen in your area costs $0.02/lb. It would be a good idea to test the sensitivity of the optimumplant design and economics to uncertainty in the selling prices of the product and the raw material.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate, animportant consideration with oxygen and hydrocarbons. Remember that you will be there for thestart-up and will have to live with whatever design decisions you have made.

Reference

World Patent Applications 00/14035 and 00/14180 to Dow.

A-II.1.4 Butadiene to n-Butyraldehyde and n-Butanol

(Bruce Vrana, DuPont, January 2000)

n-Butyraldehyde is conventionally produced from propylene and highly toxic synthesis gas in theso-called oxo process. The n-butyraldehyde is used to make 2-ethyl hexanol via aldol condensationas well as n-butanol. These oxo alcohols are frequently used, in either the alcohols or ester form, assolvents.

Because propylene is frequently quite expensive and in short supply, BASF has applied for a patenton a new route to n-butyraldehyde and/or n-butanol starting from butadiene. They found that ahomogeneous palladium acetonylacetonate catalyst with phosphine ligands would allow butadieneto react with n-butanol to produce 1-n-butoxy-2-butene (nBB). nBB will then react with more n-butanol to produce the acetal, using a homogeneous phosphine modified ruthenium catalyst. Theacetal can be hydrolyzed to n-butyraldehyde, or hydrogenated and hydrolyzed to n-butanol usingthe same Ru catalyst.

CH2=CHCH=CH2 + BuOH → BuO-CH2CH=CHCH3 [nBB]

nBB + BuOH → (BuO)2CHCH2CH3 [Acetal]

Acetal + H2O → O=CHCH2CH2CH3 + 2 BuOH

Acetal + H2 + H2O → 3 BuOH

CD-A-II-12Unfortunately, in the first reactor, a side reaction produces 2-butoxy-3-butene (iBB). The iBB canbe isomerized to nBB using an acid ion exchange resin or a Pd catalyst. Unfortunately, thisisomerization reaction is likely to be equilibrium limited.

BASF also found that while this reaction works well with pure butadiene, it will also work with"crude" butadiene, the C4 olefin cut from an ethylene cracker. The butenes in the C4 cut are inertunder the reaction conditions.

Your company has asked your group to determine whether this new technology should be used inyour Gulf Coast plant, and if so, what the economic optimum feedstock and product would be. Thegoal is to maximize the net present value (NPV) of the project. Based on past experience, youknow that you will have to be able to defend any decisions you have made throughout the design,and the best defense is economic justification.

Your company has 200 MM lb/yr of crude butadiene, which is currently being burned for fuelvalue. Thus, one possible feedstock would be the butadiene contained in the crude. You wouldreceive a credit for the unused C4's in the stream, so you would only have to pay fuel value for thebutadiene you actually consume in the process. Of course, the inert C4's will dilute the reactorcontents, making it larger, and complicate the separation train. As an alternative, you couldpurchase pure butadiene for $0.15/lb in 2001 dollars, which would result in smaller vessels.

The composition of your plant's C4 cut, which has already passed through your MTBE plant toreact away the isobutylene, is:

For a product, you could produce n-butyraldehyde or n-butanol, or some combination of the two.Your marketing organization believes they could sell the aldehyde for $0.40/lb, and n-butanol for$0.40/lb also, both in 2001 dollars.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate.

World Patent Application 98/41494 to BASF

Methyl methacrylate (MMA) is a monomer or comonomer in many polymers, most notably

Plexiglas (R). Although it is the methyl ester of methacrylic acid, it is not often produced frommethacrylic acid.

BASF has recently patented a reactive azeotropic distillation process to produce esters frommethacrylic acid and alcohols, involving a total of 3 columns. Although the patent example is forbutyl methacrylate, they claim methyl methacrylate as well.

Design a process and plant to produce 100 MM lb/yr of MMA from methacrylic acid that yourplant already produces. Use the process concept that BASF introduces, with appropriatemodifications (improvements) for MMA.

Your process design must be supported by rigorous distillation simulations. VLE and LLE data areavailable in the DECHEMA Chemistry Data Series (Gmehling et al., 1980). Do not blindly useactivity coefficients from a simulation program.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate.Assume a U.S. Gulf Coast location on the same site as a large oil and petrochemical plant. 99.95%pure MMA can be sold or transferred for $0.60/lb, according to your marketing organization. Theacid feed contains 5% water (by weight). Because it is impure, the cost of the acid in the stream is$0.40/lb. Your marketing organization projects that the long-term average price of methanol is$0.40/gal.

U.S. Patent 5,734,074 to BASF

A-II.1.6 Coproduction of Ethylene and Acetic Acid from Ethane

(Bruce Vrana, DuPont, January 2000)

Ethylene is the largest-volume organic chemical, with world production over 50 billion pounds peryear. It is normally produced by steam cracking of ethane or heavier hydrocarbons. Acetic acid isanother large-volume chemical, with annual world production in the billions of pounds. Aceticacid is normally produced using the Monsanto process from methanol and highly-toxic carbonmonoxide, although there are some older technology plants still running.

CD-A-II-14Saudi Basic Industries (Sabic) has applied for a patent on a new catalyst which will coproduceethylene and acetic acid from ethane and air. Their catalyst is a phosphorus-modifiedmolybdenum-niobium vanadate. At different phosphorus levels, the catalyst will produce differentratios of ethylene to acetic acid. Selectivity to the two products is also a function of conversion(i.e., space velocity). As conversion increases, the selectivity to ethylene decreases and theselectivity to acetic acid increases. However, the total selectivity to the useful products decreasesas conversion increases. The process runs at higher pressures, about 200 psig, than a conventionalethylene furnace.

Your company manufactures 2 MMM lb/yr of ethane which is currently being produced at yourGulf Coast plant and sold for $0.07/lb in 1999. Your team has been asked to evaluate the economicviability of the Sabic process for your plant, as a way of upgrading your product and increasingyour sales revenue. Your job is to determine the economic optimum design, producing whateverproducts will maximize the net present value (NPV) of the project. You may consume all or part ofthe ethane which is available and make any ratio of ethylene to acetic acid which can be producedby the catalyst. Based on past experience, you know that you will have to defend any decisions youhave made throughout the design, and the best defense is economic justification.

Your marketing organization believes they can sell ethylene for $0.25/lb in 2000 dollars. Althoughthey are less certain because it is a new product for your company, they also believe they can sellacetic acid for $0.19/lb in 2000 dollars. It would be a good idea to test the sensitivity of theoptimum plant design and economics to uncertainty in the selling prices of both products.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate.

Reference

World Patent Application 99/13980 to Sabic

A-II.1.7 Methylmethacrylate from Propyne

(Bruce Vrana, DuPont, January 1999)

Methyl methacrylate (MMA) is a monomer or comonomer in many polymers, most notably

Plexiglas (R). The conventional process has many drawbacks, including the use of sulfuric acid as acatalyst. Most manufacturers neutralize the sulfuric acid with ammonia, producing byproductammonium sulfate which must be sold or disposed of. HCN is also used in the process, requiringthe MMA plant to be linked to a source of hazardous HCN.

Shell has patented a new process with several advantages over conventional MMA processes. Amajor advantage is that neither HCN nor sulfuric acid are used. Shell found that propyne can becarbomethoxylated (reacted with CO and methanol) to produce MMA directly. The maindisadvantage is that propyne is not normally considered a viable feedstock due to its scarcity and

CD-A-II-15the impurities it contains. Shell's new catalyst tolerates impurities in the propyne much better thanprior catalysts.

Your job is to develop a scenario for Shell to commercialize this process. You must first find asuitable feedstock for this process from the normal refinery and/or petrochemical streams available.Producing propyne to provide the feedstock is discouraged, due to high cost. Having found astream which contains suitable quantities of propyne in high enough purity for this process to befeasible, design a plant to produce 100 MM lb/yr of MMA by the new Shell process. Determine theoverall economic feasibility of the plant.

The plant design should be as environmentally friendly as possible. Recover and recycle processchemicals to the maximum economic extent. Also, energy consumption should be minimized, tothe extent economically justified. The plant design must also be controllable and safe to operate.

Assume a U.S. Gulf Coast location on the same site as a large oil and petrochemical plant. MMAcan be sold or transferred for $0.60/lb, according to your marketing organization. Value thepropyne as appropriate for alternative uses for the stream (i.e., if the stream you are using isnormally burned, value the propyne at fuel value). A major gas vendor is willing to locate acrossthe fence from you and supply CO at the required pressure for $0.12/lb. Your marketingorganization projects that the long-term average price of methanol is $0.40/gal.

Reference

U.S. Patent 5,719,313 to Shell Oil Company

A-II.1.8 Mixed-C4 Byproduct Upgrade

(Leonard A. Fabiano and Robert Nedwick, Lyondell, January 1999)

Your company is a major player in commodity petrochemicals, specifically producing olefins viathe cracking of ethane, propane, butane and naphthas. At one of your Gulf Coast sites, the majorproducts are ethylene and propylene in addition to a number of smaller fuel streams. The crude C4product, which because of the feed mix has been a relatively small portion of the product slate, iscurrently being sold at fuel value. Now, due to a change in feed mix, the C4 yield from thecracking furnaces has increased significantly. Management would like to upgrade this streamabove fuel value. The expected feed composition and flow rate are as follows:

Flow rate, lb/hr 100,000

The company would like to maintain its focus on commodity chemicals and is interested in highvolume products. Your project team has been assembled to determine:

1. What components are worth considering for recovery?

2. What processing options are available for the components of interest?

3. What is the most economical processing route?

and to develop a design package that will meet a 15% return on investment.

A-II.1.9 Hydrogen Peroxide Manufacture

(Bruce M. Vrana, DuPont, January 1999)

Hydrogen peroxide is an oxidant used in many markets, including the pulp and paper industry.Almost all of the world capacity is based on alternately hydrogenating and oxidizing an expensivealkylanthraquinone.

Enichem has applied for a patent on a process based on oxidizing carbon monoxide in a complexaqueous solution. Rather than using expensive hydrogen, this process incorporates the hydrogenfrom water. The overall chemistry is:

CO + H2O + O2 → H2O2 + CO2

The application cites data with reactor productivities comparable to or even better than theconventional chemistry. Design a process and plant to produce 100 MM lb/yr of 50% H2O2 usingthis proposed reaction path.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be controllable and safe to operate.

Assume a U.S. Gulf Coast location on a large plant complex. H2O2 can be sold or transferred for$0.60/lb, according to your marketing organization, on a 100% basis. A major gas vendor is willing

CD-A-II-17to locate across the fence from you and supply CO at the required pressure for $0.12/lb and oxygenfor $0.02/lb.

Reference

European Patent Application 808796 by Enichem.

A-II.1.10 Di-tertiary-butyl-peroxide Manufacture

(Leonard A. Fabiano, ARCO Chemical, January 1995)It is desired to design a process to produce 100 million pounds per year di-tertiary-butyl-peroxide(DTBP) based primarily on a Texaco patent. DTBP is an important chemical that has use, forexample, as a catalyst in various organic syntheses and has special utility as an additive to dieselfuel formulations to improve its combustion characteristics. It behaves in an analogous way todiesel fuel as octane enhancers (e.g., MTBE) behave in gasoline (see U.S. Patent 5,312,998,column 1, lines 29-33). The product must contain less than 0.3 weight percent tertiary-butyl-alcohol (TBA) and essentially no other peroxides. The plant will be constructed at a Gulf Coastlocation adjacent to a feedstock-producing facility. Texaco and ARCO have facilities in this area.

Specific kinetic data are not available but hourly space velocities are provided in the Texaco patent(80-100˚C, 1-2 vol. TBHP per vol. catalyst per hour – U.S. Patent 5,345,009, column 4, lines 23-44). Phase equilibrium data are to be developed from the DIPPR databank and UNIFAC estimatesusing ASPEN PLUS.

Specifics

Your group is requested to develop and analyze a process to produce DTBP based on informationprovided in U.S. Patent 5,345,009 assigned to Texaco Chemical Company, and U.S. Patents5,288,919 and 5,312,998 assigned to ARCO Chemical Company.

Assistance will be provided in making decisions, but will be very specific with references in theopen patent literature. It should be apparent in this problem statement of this most timely processstudy that I must be careful not to release proprietary information which is contained in a veryrecent patent application for which I am one of the inventors. The results of this comparison of theTexaco process, as devised by your group, with the ARCO process is typical of an exercise that allcompanies must undertake to analyze the economic viability of all new ventures.

We are interested in comparing the Texaco technology with the confidential process developed byARCO. However, you are expected to be very creative and devise a continuous process tominimize costs. It is suggested that you focus on Texaco patent (5,345,009 - column 2, lines 65 tothe end, and column 3, lines 1-6). Paraphrasing, di-tertiary-butyl-peroxide (DTBP) is formed whentertiary-butyl-hydroperoxide (TBHP) and an enhanced amount of tertiary-butyl-alcohol (TBA) arebrought into contact with a palladium-coated, carbon catalyst; that is,

TBA + TBHP → DTBP

CH3 CH3 CH3 CH3

TBA + Isobutylene ( iC 4= ) + TBHP → DTBP

The two routes above are basically the same since TBA under the proper conditions and in thepresence of a catalyst reacts to form isobutylene and water according to the reversible reaction:

TBA = iC 4= + H2O

The isobutylene is the molecule that reacts directly with the TBHP.

ARCO Patent 5,312,998 (column 3, lines 31-40) offers the same possibilities. TBHP iscatalytically reacted with TBA to form DTBP. Isobutylene can be added to the reaction mixtureand it is generally advantageous to use a substantial excess of TBA and/or iC4 relative to the TBHPto achieve high TBHP conversion; e.g., 90% or more. Conditions for the reaction (with differentcatalysts) are proposed in U.S. Patent 5,345,009 (column 4, lines 24-33). The reaction may beconducted at a temperature within the range of about 40˚C to about 160˚C at super-atmosphericpressures. A contact time of about 0.5-10 hours is required. U.S. Patent 5,288,919 (column 2, lines19-29) suggests temperatures ranging from 20-150˚C at a sufficient pressure to ensure a liquid-phase reaction.

U.S. Patent 5,312,998 (column 2, lines 5-19) suggests that there can be a two-liquid phase reactioncarried out in the temperature range of about 70-110˚C.

Similarities - Despite the Differences in Catalysts

U.S. Patent 5,345,009 (column 3, lines 15-29) suggests a typical feed stock for the Texaco process,but specifics of other components are not described. U.S. Patent 5,288,919 (column 3, lines 29-39)suggests a typical debutanized feed stock composition of 58 weight % TBA and 40 weight %TBHP, with the remainder comprised of 0.2% methanol, 1.3% acetone, and 0.5% water. For theTexaco process, let's use a mixture of 70% TBA, 30% TBHP and assume that this mixture makesup 98% of the mixture based upon the ARCO patent. The remaining 2% is assumed to be as above.

Note that TBA and DTBP, as well as TBA and water, form azeotropes.

Let's brainstorm and develop several likely candidate processes to evaluate and perhaps comparebefore we embark on detailed evaluations.

Alternative Process

ARCO produces TBHP-70, a possible "purified" feedstock for the reaction:

TBHP + iC 4= → DTBP + TBA

TBHP-70 is essentially 70% TBHP and 30% water. Would this provide an economically viableprocess?

References

U.S. Patent 5,345,009 (September 6, 1994).

Our company, BCI (Better Chemicals Inc.) has recently discovered a new product which we intendto manufacture in the near future. This product uses vinyl acetate as one of the main raw materials.We expect to use 300 MM PPY of vinyl acetate in our new process. In reviewing the economics ofour new product, we found that it was negatively impacted by the relatively high market price ofvinyl acetate ($0.44/lb). A closer investigation showed us that the most popular route to vinylacetate is from ethylene and acetic acid oxidized by oxygen. The site where our new process willbe constructed happens to use all three ingredients needed for vinyl acetate. Very favorable, longterm contracts for their use have been negotiated. We thus find that we can obtain large quantitiesof acetic acid for $0.27/lb and ethylene for $0.20/lb. Oxygen costs us $0.02/lb. With these rawmaterial prices, we feel that we can manufacture vinyl acetate far below the market price of$0.44/lb and thus make our new product that much more profitable. In assessing the project tomanufacture our own vinyl acetate, we used some approximate estimating techniques [1] toevaluate the investors rate of return we could expect from a 300 MM PPY vinyl acetate plant as afunction of the onsite capital investment. In these calculations, the onsite cost consists of theinstalled cost of all process equipment within battery limits. We estimate the offsite cost to be 45%of the onsite cost and apply a 25% contingency such that the fixed capital is related to the onsitecost as

Fixed Capital = 1.25 (onsite + 0.45 onsite)

The results of our venture guidance calculations are shown in the figure below

CD-A-II-20While we do not know exactly how much we need to invest into the vinyl acetate process (this isone of the questions we have for you), we crudely estimate it to be less than $50-60 MM onsite.Since the cost of capital is 12%, we therefore expect this to be a profitable venture.

We now turn to the technology of the vinyl acetate process. Reference [2] gives an overview of theprocess and states that the main reaction is

H2C = CH2 + CH3COOH + 1/2O2 → H2C = CHOOC-CH3 + H2O (R1)

Reference [2] also indicates that the most economic route to vinyl acetate, when acetic acid isavailable, is to convert the raw materials to product in the vapor phase over a palladium catalyst.We therefore asked our research chemists to develop a catalyst suitable for the operation. Theyfound a suitable catalyst by impregnating a silica base with 2% palladium along with some otherproprietary chemicals. The chemists performed numerous experiments with the catalyst and foundthat it is quite selective towards vinyl acetate and quite active as measured in its space time yield(STY, grams of vinyl acetate/hr per liter of catalyst). The only significant side reaction we couldnotice is the combustion of ethylene to carbon dioxide and water

H2C = CH2 + 3O2 → 2CO2 + 2 H2O (R2)

Once the catalyst was developed our chemical engineers designed a kinetic study using alaboratory-scale reactor to quantify the performance of the catalyst for the purpose of designing acommercial-scale reactor. For commercial purposes the catalyst support will be pelletized such thatthe bulk average density of the final catalyst is 30 lb/ft 3. The following rate expressions wereobtained:

O2 H 2O (1 + 6.8 p HAc )  min ⋅ lb catalyst 

and for R2:

r2 = 1.9365 ⋅10 5 e −10,116 / T

In these expressions, T is absolute temperature in kelvins and p is the partial pressure of a

component in psia. We also calculated the heat of reaction in the ideal gas standard state (25°C, 1atm) by using available heats of formation of the components. The standard state heat of reaction is-42.1 kcal /mol of vinyl acetate for R1 and -316 kcal /mol of ethylene for R2. The reactions arethus quite exothermic, which we also observed in the laboratory.

Based on this information BCI is requesting that your company design a cost effective process tomake 300 MM PPY of crude vinyl acetate. Since vinyl acetate and water form a heterogeneousazeotrope we refer to crude vinyl acetate as the acetic acid “free”, liquid product which could be

CD-A-II-21decanted off from the reaction water. The crude vinyl acetate will then contain water up to itssolubility limit at say 20°C which is about 5 mol% water. The acetic acid in the crude vinyl acetatemust be less than 0.1 mol%. BCI has existing columns on site capable of removing the remainingwater, acetic acid and other byproducts from the crude vinyl acetate. We also suggest that youwould use one of many standard principles (e.g. carbonate wash) for removing the byproductcarbon dioxide from the reaction mixture. In your flowsheet you need not design or analyze thecarbon dioxide removal step in detail but simply assume that 99.5% of the carbon dioxide will beselectively removed from any stream sent to such a facility. The size (and cost) of the carbondioxide removal unit will be proportional to the flow rate and composition of the stream sent to it.You may cost estimate the carbon dioxide removal unit as two packed towers (one absorber andone desorber) each with 30 equivalent stages. In the first tower, the absorber, CO2 is absorbed in acold liquid (assume water) containing a carbonate. In the second tower, the desorber, the CO2 isliberated by reboiling the recirculating liquid. Based on our requirement that the desorber mustoperate at atmospheric pressure and that we would like to use cooling water for the absorptioncooler, we have estimated the following heat load requirements for the CO2 removal unit. Thisshould aid you in estimating the diameter of the towers and the sizes of the heat exchangersdepending on the nature of the stream you opt to purify.

Mol% Carbon Dioxide in the Vapor Heating Requirement in the

You may further assume that acetic acid is available from our tank farm as a liquid at 30°C. Youmay also assume that both ethylene and oxygen are available from separate gas headers at 200 psigand 30°C. The ethylene gas is 99.9% pure, the balance being ethane. The following utilities andservices are available as needed at the battery limits. Costs are in 1996 dollars

150 psig steam $5/1,000 lb

In designing the process we would like you to propose a design which minimizes the total productcost of crude vinyl acetate at the nominal rate of 300 MM PPY of pure vinyl acetate. Assume a

CD-A-II-2290% operating utility (7,884 hr/yr) and assume that 99% of the vinyl acetate in the crude stream canbe recovered. The results we expect from your work include

• An optimized flowsheet• Total installed equipment costs (onsite cost)• A profitability analysis of the project• A control scheme based on an in-depth operability analysis of the process

Physical properties for all components required in this study should be readily accessible frompublicly available sources (e.g. DIPPR, HYSYS.Plant, etc.). This also pertains to mixtureproperties with the possible exception of the vinyl acetate (1)/water (2) binary. We thereforeprovide you with our best estimate of the VLE and LLE data for this pair.

A-II.1.12 PM Acetate Manufacture

(Leonard A. Fabiano, ARCO Chemical, January 1993)

PM Acetate (propylene glycol mono-methyl-ether acetate) is a specialty solvent used in resins,

coatings and cleaner formulations. Current sales volumes are 10 MM lb/yr and it is being producedbatchwise by outside "tollers". Due to expected increases in demand, the PENNCO (yourcompany's name) is interested in building its own continuous plant in the Houston area. Theeconomic size must be determined that will yield a 15% after tax return while the sales build to 20MM lb/yr in three years. Consider first a 20 MMlb/yr facility which will be integrated into anexisting facility. Our R&D groups have developed a considerable amount of data on the process;i.e., chemical kinetics and VLE data. This information will be supplied after the design group signsa non-disclosure agreement with ARCO.

PG + HOAc PG Acetate and PG Diacetate + H2 O

MeOH + HOAc MeAc + H2 O

Data have been developed on a boiling reactor concept that utilizes a liquid catalyst and a fixed-bedreactor concept that utilizes an acid resin catalyst. The fixed-bed option offers several advantages,in particular, in raw materials cost and handling, and in materials of construction. It is requestedthat you investigate the fixed-bed concept and compare it with a reactive distillation concept thatutilizes the solid catalyst.

The expected market price, chemical kinetics and VLE data, and utility costs will be supplied at alater date. Where VLE data are lacking you may use the UNIFAC correlation. Your company hasaccess to ASPEN PLUS which has a reactive distillation subroutine (RADFRAC).

CD-A-II-24 Unreacted PM, HOAc

PM Waste Water Reaction Separation PMA Product HOAc

Heavies Waste

Figure 1. Simplified Flowsheet for Fixed-bed Process

A-II.1.13 Propoxylated Ethylenediamine

(Brian E. Farrell and David D. Brengel, Air Products and Chemicals, January 1994)

Ethylenediamine (EDA) is a versatile building block in the chemical industry for amine-basedcompounds.

NH 2 H2 N EDA

A family of amine compounds can be formed from the reaction of EDA with propylene oxide (PO).

CH3 PO

Between 1 and 4 moles of PO can be added per mole of EDA. The monopropoxylated EDA can beused as an intermediate in the synthesis of a polyurethane catalyst. The di- and tri-propoxylates canbe used as cross-linkers for epoxy systems. The fully propoxylated molecule is used as a cross-linker in polyurethane systems.Your assignment, should you decide to accept it, is to synthesize and purify each of the EDA-POreaction products. The required amount of each product will be determined according to marketdemand. IMF, the company that you work for, has performed extensive market research and willprovide you with an estimate of market demand and selling price for each of the four compounds.The IMF research department has synthesized the four materials in small quantities and will makeavailable their findings with regard to reaction kinetics and thermodynamics. You will beresponsible for designing a reactor system and distillation process that best meets the anticipatedmarket demands, while simultaneously maximizing IMF's profits.

CD-A-II-25A-II.2 PETROLEUM PRODUCTS

A-II.2.1 Fuel Additives for Cleaner Emissions

(E. Robert Becker, Environex, January 1993)

Carbon monoxide and ozone levels are in excess of the National Ambient Air Quality Standards inthe Northeastern states, which constitute a corridor from Virginia to New England. The principalsource of carbon monoxide are emissions from automobiles. The coalition of Northeasternregulators have mandated cleaner burning fuels for the region; however, demand is uncertain sincethe member states can opt into the plan until 1995. The use of methyl-tert-butyl-ether (MTBE) asan octane enhancer provides significant reductions in carbon monoxide emissions.

Your company has technology for the production of MTBE. Your assignment is to providemanagement with a cost estimate for a 100,000 gallon per day MTBE plant in the Philadelphia tri-state area. Your report should estimate the product prices necessary for annual production rates of100,000, 70,000, and 50,000 gallons. You have a stream of butane available from an adjoiningrefinery and you have to purchase methanol from a nearby chemical plant. Steam can be purchasedfrom a cogenerator.

The process involves the dehydrogenation of isobutane to isobutene which is reacted with methanolto produce MTBE. Particular attention should be given to the dehydrogenation reactor design andoperation. Technical and economic data for the design are attached.

The catalyst is 0.3 cm chromia alumina spheres with 0.48 void fraction and 1,200 kg/m3 bulkdensity. The carbon is removed from the catalyst by burning in air at a rate of 0.1 kg carbon/kgcatalyst-hr. The maximum catalyst temperature is 740˚C. The catalyst is replaced annually.

The reaction of isobutylene and methanol is assumed to go to 98% equilibrium without sidereactions. The dehydrogenation reaction produces isobutene, hydrogen, propylene, and methane.

Cost and Economic Data

95% isobutane-5% n-butane is $ 0.70/gallon

Methanol is $0.75/gallonSteam at 700˚C and 10 bar is available at $8.00/1,000 lbElectricity cost is $ 0.07/kWhrFuel gas is valued at $2.00/MMBtuCooling water is $0.15/1,000 galCatalyst is $15/kg

Annual effective interest rate = 12% per year

Project life 10 yearsMinimum investor’s rate of return (IRR) is 15%

A-II.3 GAS MANUFACTURE

A-II.3.1 Nitrogen Rejection Unit (from natural gas)

(William B. Retallick, Consultant, January 2002)

This unit is part of a gas plant, which prepares raw natural gas for sale to a pipeline. The front endof the gas plant has already removed the natural gas liquids from the gas. It remains for therejection unit to remove nitrogen and also recover helium, a valuable by-product. Flow diagramsfor the unit are included in a paper by Scott Troutmann, of Air Products and Chemicals, and KimJanzen, of Pioneer Natural Resources. The unit uses two stripping columns. You can produce aside stream from the first stripping column that contains about 50 mol% nitrogen. This will be usedto fuel the gas turbines, which drive the compressors.

The feed consists of two streams:

1. Pipeline gas is to be delivered at 1,200 psig, containing no more than 2 mol% N2.

2. Crude helium product contains at least 65 mol% helium, a maximum of 1 mol% methane, with the balance N2, and is delivered at 1,200 psig. Recovery of helium is at least 96 mol%.

3. The selling price of crude helium is $25 per 1,000 ft3 of helium content.

4. When heat is transferred (irreversibly) with a temperature difference, ∆T, the lost work is Q∆T/T, where T is the temperature of the warm fluid.

At cryogenic temperatures, where T is smaller, the losses are greater. Hence, to avoid increases in the lost work as T decreases, the minimum internal temperature difference (MITD) must be reduced. As you carefully select the MITD, consider the range of 1 - 6 K for your design.

5. Simplify your calculations with the units K, kg and atm.

6. Purchased electricity costs $0.70 per kWh.

7. The plant is located in Texas.

8. The cryogenic vessels and exchangers are of 304L stainless steel.

9. The heat exchangers are plate exchangers.

10. You can display the economics of your process by graphing the investor’s rate of return (IRR) as a function of the cost of the feed divided by the sales price of the gas.

As the semiconductor industry goes to submicron and deep submicron designs, the purityrequirement for nitrogen gas is becoming higher and higher. The current specification for nitrogenrequires the impurity levels to be below 10 parts per billion by volume.

Your company, UltraPureGas, is approached by a major semiconductor manufacturer (Advanced

SemiCon) to submit a proposal to supply 200 ton/day of nitrogen at a pressure of 10 bar absolute totheir megafab in Austin, Texas. The maximum allowable total impurities content (excluding noblegases such as argon, neon, and helium) is 10 parts per billion by volume. The customer alsoindicated that to avoid potential particulate contamination, nitrogen product compressors should beavoided. You, the lead process engineer for this project, are asked to come up with a low-costdesign (which means you have to compare the different known processes and/or invent newprocesses and find the low-cost option).

A-II.3.3 Nitrogen Production

(Rakesh Agrawal, Air Products and Chemicals, January 1999)

Our Polymers Division needs a supply of moderately high purity nitrogen for its productionapplications. We would like to study the feasibility of incorporating new nitrogen plants with aminimum capacity of 5,000 SCFH (to handle current production) with the possibility of expansionto 40,000 SCFH. This plant is projected for 2005 when we expect the polymer market to expandsignificantly.

I am writing to you at this time to request a preliminary design for a nitrogen plant that produces20,000 SCFH of polymerization grade nitrogen. In your design you will need to compute the priceof nitrogen that yields an investors rate of return (IRR) of 15%. You should compare thiscalculated price with the price given in the Chemical Marketing Reporter.

Attached are relevant data on feedstocks, product specifications, utilities and economic data thatshould be useful for this design project. Additional data are also available in several articles in thelibrary. For this production rate there are several competing technologies. To produce acompetitive design, we would like to consider all of these technologies. These articles form onlythe start of your literature search. You will need to investigate potential ideas for this projectthoroughly.

When preparing your design, you may also make the following assumptions:

1. Nitrogen product should be delivered as dry gas at ambient conditions

2. The plant should be designed for 8000 hours of operation per year 3. The product nitrogen should be at least 99% pure

Product Specifications

20,000 SCFH nitrogen gas

Minimum Nitrogen 99 vol %

Feedstock

Air at ambient conditions

CD-A-II-30Utilities

Cooling Water:

90°F supply temperature

115°F maximum return temperature

Steam System:

Saturated Steam from Offsite Boilers

Available at 150 and 600 psig

Process Water

Available at 90°F

Ambient Design Temperature:

100°F dry bulb, 90°F wet bulb

Economic Data

The following data are necessary for the economic evaluation. These include estimates needed forthe 2005 analysis and follow trends over recent years.

A-II.3.4 Krypton and Xenon from Air

(Rakesh Agrawal and Brian E. Farrell, Air Products and Chemicals, January 1991)

Krypton and Xenon are rare gases which are normally recovered from air. Recently, their demandhas been on the rise. They are used in various applications - in several medical devices, long-lasting light bulbs, nuclear magnetic resonance, etc. The concentration of each of these gases in airis extremely low (below 5ppm). This makes their recovery from air challenging and technicallyexciting.

To produce reasonable quantities of krypton and xenon, both gases are recovered from large-tonnage plants for air separation that produce oxygen in quantities greater than 500 tons/day. These

CD-A-II-32large plants are cryogenic in nature and operate at temperatures as low as -195˚C. Air is composedprimarily of oxygen (20.95 mole %), nitrogen (78.12%) and argon (0.93%). However, besidesargon, it has several contaminants such as hydrogen, helium, neon, carbon monoxide, methane andother hydrocarbons, water and carbon dioxide. Most of these contaminants are in much higherconcentrations than krypton and xenon. The feed to the cryogenic air separation unit (ASU) ispressurized to about 6 atm before water, carbon dioxide and some hydrocarbons are adsorbed onmolecular sieves. The air stream is cooled to near its dew point and distilled to recover nitrogen,argon and oxygen. Of these three constituents, nitrogen is the most volatile and oxygen the least.Contaminants such as hydrogen, helium, neon, and carbon monoxide concentrate at the top of thedistillation column and leave with the nitrogen product. Krypton and xenon, along with methane,ethane, propane and some ethylene and propylene, are concentrated in the liquid oxygen (LOX)collected at the bottom of the distillation column. All of these components have boiling pointshigher than oxygen and are heavier. The efficient and economical recovery of krypton and xenonfrom LOX is the subject of this design project.First, a conventional plant to recover krypton and xenon from the LOX will be designed. In thisprocess, a portion of the LOX stream containing krypton, xenon and other hydrocarbons iswithdrawn from the bottom of the main distillation column and passed through a bed to adsorb allthe heavier hydrocarbons, including propylene and ethylene. None of the methane is adsorbedwhile some of ethane and propane are adsorbed. The LOX stream is fed to the top of the firstdistillation column to concentrate krypton and xenon (since the concentration in the feed LOX isbelow 50 ppm). However, the concentration of krypton and xenon in the bottom distillate from thiscolumn cannot be increased by more than a factor of about ten. The primary reason is that, alongwith krypton and xenon, hydrocarbons concentrate in the liquid phase. Concentrations of methanein liquid oxygen exceeding 50 ppm are unacceptable because they are explosive and present asafety hazard. The vapor from the top of this column is returned to the main distillation columnand the liquid oxygen from the bottom, containing krypton, xenon, methane, ethane and propane, isvaporized in heat exchangers. The vaporized stream is heated to about 550˚C and sent to a catalyticunit to burn the hydrocarbons. The effluent from the catalytic unit is cooled and is passed througha molecular sieve adsorbent to remove the water and carbon dioxide formed during the reaction.The resulting stream is cooled to cryogenic temperatures, liquified and distilled to recover kryptonand xenon. The oxygen stream from this distillation step is recycled to the first distillation columnto recover krypton and xenon.After the conventional process is designed, more recent technology will be considered. It may bepossible to reject methane from the first distillation column and concentrate krypton and xenon byseveral orders of magnitude (as compared to a factor of about ten). Also, these processes can bemade inherently safe by feeding nitrogen to the stripping section of a second distillation column,thereby displacing most of the oxygen from the krypton and xenon in the stripping section. Designof these processes should expose the opportunities for integrating the krypton/xenon distillationcolumns with heat and mass from the main air distillation units.

A-II.3.5 Ultra-High-Purity Oxygen

(Mark R. Pillarella and Rakesh Agrawal, Air Products and Chemicals, January 1992)

Computers have revolutionized industry and technology over the past 15 years and can be expectedto continue to do so. Improvements in computer technology are driven by improvements insemiconductor technology. For the production of high quality, defect-free semiconductors, ultra-high purity (UHP) oxygen is essential in the etching process. Typical cryogenic processes canproduce oxygen with parts-per-million by volume impurities, but semiconductor manufacturingrequires oxygen with impurities less than parts-per-billion by volume.

CD-A-II-33Your company, OxyPure, is submitting a proposal for a multi-million dollar contract to supplyultra-high purity oxygen to a major semiconductor manufacturer (SemiCon) in Southern California.OxyPure operates a conventional oxygen plant in Southern California which produces 400 metrictons per day of 1.3 bara standard grade gaseous oxygen (99.5% oxygen, 0.5% argon, 10 ppmmethane, 0.5 ppm other hydrocarbons, 5 ppm krypton, 0.4 ppm xenon, 0.1 ppm nitrous oxide,essentially no nitrogen). The process flow diagram is shown in the Figure 2. Your processengineering team has been assigned the task of evaluating several process schemes for modifyingthe existing plant to supply the semiconductor customer.SemiCon requires 10-40 metric tons per day of 1.0 bara gaseous UHP oxygen. They haverequested that proposals be submitted for two purity specifications;(1) Less than 25 ppb of hydrocarbons; concentration levels of the other impurities acceptable.(2) Less than 5 ppb argon and less than 5 ppb of the remaining impurities.The process schemes to be evaluated are:(A) Part of the standard grade oxygen can be reacted over a palladium or another suitable noble metal catalyst at 500°C, converting the hydrocarbons and some of the oxygen to carbon dioxide and water: CH4 + 2O2 → CO2 + 2H2O The reactor effluent is passed through an adsorption bed (containing 5A or 13X molecular sieve adsorbent) to remove the CO2 and H2O (Giacobbe, 1989, 1991).(B) Part of the standard grade oxygen can be fed to a standard three-component distillation process (requiring two additional distillation columns) to remove both the light and heavy impurities (King, 1980).(C) A side stream can be withdrawn from the upper column and fed to an additional distillation column which removes the remaining impurities to produce UHP oxygen.Develop each process scheme and compare the product purity, efficiency, and economics.Necessary process information will be supplied for the conventional oxygen plant.ReferencesGiacobbe, F.W., "Use of Physical Adsorption to Facilitate the Production of High Purity Oxygen",Gas Separation & Purification, Vol. 3, 1989.Giacobbe, F.W., "Adsorption of Very Low Level Carbon Dioxide Impurities in Oxygen on a 13XMolecular Sieve", Gas Separation & Purification, Vol. 5, 1991.King, C.J., Separation Processes, 2nd ed., McGraw-Hill, New York, 1980.

CD-A-II-34 Figure 2. A conventional process for oxygen production.

A-II.4 FOODS

A-II.4.1 Monosodium Glutamate

(Robert M. Busche, Bio-en-gene-er Associates, January 1991)

In its efforts to expand into new specialty chemical markets, your company is consideringmanufacturing the flavor enhancer MSG (monosodium-L-glutamate monohydride) for the U.S.market by way of a joint venture with the Ajinomoto Company. Ajinomoto is the Japanesecompany that presently dominates the world market for MSG. The market situation in 1984 inmillions of annual pounds was:

Production Consumption

S.E. Asia 397 300

CD-A-II-35With the help of Ajinomoto, the Marketing Department believes that it can capture a 50 millionpound share of the North American market by the year 2000. Sales are expected to start at 20million annual pounds in 1992; 30 in 1994; 41 in 1996; and 48 in 1998.

Presumably, the plant design will be based on the Ajinomoto batch fermentation process convertedto a continuous mode using the aerobic bacterium Brevibacterium ammoniagenes. However, yourResearch Department recently was able to isolate a gene for a hemoglobin-like molecule from theaerobe Vitreocilla and express it in Brevibacterium. The recombinant cells contain hem andactive hemoglobin. As a result, they appear to grow faster and to considerably higher cell densitiesthan the conventional cells, especially when dissolved oxygen is less than 5% of air saturation.

Before committing to the joint venture, your president would like you, as Director of the CorporatePlanning Department, to assess the expected economic performance of the Japanese process, asoperated at your plant in Iowa and also to ascertain the sensitivity of the process economics to theuse of the new organism.

The Japanese process operates with two fermenter stages. In the first stage, cells are grown to adensity of 17.5 g/liter before inducing product expression. The cells are grown from glucose (cornsyrup) according to the overall reaction:

C6H12O6 + 3O2 = 3CH2O + 3CO2 + 3H2O

Six hours are allowed for growth.

The product is produced form the resting cells in the second stage, at pH 7.0-8.0, over a 28 hourperiod, at a concentration of 90 g/liter. The overall reaction to products is:

C6H12O6 + 2.2065O2 + 0.843NH3 = 0.843 C5H9O4N + 1.785CO2 + 3.471 H2O

Glucose conversion is essentially 100%. There is reason to believe that, with the new aerobe,production time might be reduced and cell density increased to, hopefully, 50 g/liter and, perhaps100 g/liter. The allowable cell density will depend on viscosity restrictions to aeration performanceof the new bacterium.

A-II.4.2 Polysaccharides from Microalgae

(Robert M. Busche, Bio-en-gene-er Associates, January 1986)

Research Department has discovered a way to produce polysaccharides (also known as water-soluble gums or biopolymers) from Porphyridium cruentum, a marine microalga. Process data areprovided in the reference. The product is expected to find uses in existing food markets as a water-binding thickening agent, competing with such products as xanthan gum, agaur, alginates andcarboxymethylcellulose. A very large potential new use is for enhanced oil recovery, where it canbe used to increase the viscosity of sweep water relative to that of crude oil so as to promote themobility of the residual oil in the reservoir. In this service, biopolymers are injected at a rate of 1.4to 1.7 lb/barrel of oil recovered. Excluding the polymer, the cost of the polymer/sulfonatesurfactant flood amounts to $30 to $40 per barrel of oil (including capital charges).

Your management has asked you to determine if the new product can be produced at a low enoughcost to compete in the food and/or EOR markets.

Reference

Anderson, D.B., and D.E. Eakin, A Process for the Production of Polysaccharides fromMicroalgae, Battelle Pacific Northwest Laboratories, Richland, WA (1985).

A-II.4.3 Alitame Sweetener

(Robert M. Busche, Bio-en-gene-er Associates, January 1987)

A new sweetener, named Alitame by its inventors in your Research Division, is a dipeptide amideof L-aspartic acid and D-alanine. In contrast, aspartame, the amino acid-based sweetener currentlyapproved by the FDA, is a dipeptide ester and contains L-phenylalanine instead of D-alanine. TheNew Products Department has tested the new material in a variety of uses and claims that it isstable enough for use in baked goods and has a longer shelf life than aspartame. It is also 12 timesas sweet as aspartame and would not be harmful to people with the metabolic disorder,phenylketonuria, who must limit the intake of substances containing phenylalanine. Use isprojected in foods, beverages, toiletries, and pharmaceuticals.

Alitame is made in a patented process from the corresponding acid and amine. Although alaninecan be purchased from the Japanese, your company is interested in producing both precursors if

CD-A-II-37economically attractive. You have been asked to evaluate the possibilities and recommend a courseof action that is economically viable.

A-II.5.1 Generic Recombinant Human Tissue Plasminogen Activator (tPA)

(Scott L. Diamond, University of Pennsylvania, January 2000)

Setting:

Plasminogen activators are powerful enzymes that trigger the proteolytic degradation of blood clotsthat cause strokes and heart attacks. Genentech owns the patent for tPA, and currently sells 100 mgdoses of recombinant tPA (activase) for about $2,000. The annual sales for tPA are about $300MM/yr. However, the patent for tPA will be expiring soon. In response, Genentech has developeda next generation, FDA-approved, plasminogen activator called “TNK-tPA” which is slightly easierand safer for clinicians to use.

While a generic form of tPA may not compete well against TNK-tPA in the U.S., there may existthe opportunity to market a low-cost generic tPA in foreign markets where urokinase andstreptokinase are low-cost (~$200/dose) alternatives that are associated with increased bleedingrisks. Additionally, reduced healthcare reimbursements to U.S. hospitals may allow a generic tPAto compete against TNK-tPA or activase.

Process:

Produce recombinant tPA using CHO cells. Since Genentech will not license their CHO cells, yourgroup will be responsible for cloning the human tPA gene and creating a stably expressing cell linefor your process.

Constrants:

1) The product must be sold as a lyophilized, sterile powder (100 mg/bottle).

5) Your annual production will need to range from 30 to 100 kg/yr.

2) Design a reverse osmosis/deionized water purification system to supply all process water.

3) Determine the steam requirements for sterilization of the bioreactors.

4) Does an economic opportunity exist for the production of generic tPA? Assume that Genentech is your only competitor.

5) Estimate the actual production cost per 100 mg/dose for Genentech to make tPA.

Assumptions:

1) Your reactor will use serum-free growth medium.

2) You have licensed the use of a hybridoma cell line that secretes tPA monoclonal antibody for the development of your affinity columns (life of column is 3 years). The license costs $120,000/yr.

Prerequisite:

The members of this design group must have completed ChE 479, Intro. to Biotech. and Biochem.Eng., or the equivalent.

A-II.5.2 Penicillin Manufacture

(Robert M. Busche, Bio-en-gene-er Associates, January 1990)

Your large pharmaceutical company controls a major share of the worldwide penicillin market,which in 1985 reached about $600 million. However, your plants are relatively old and completelydepreciated, with rising production costs. Management is alarmed that over recent years somemarket share has been lost to companies entering the market with new plants. A decision must bemade as to whether to milk the present business as a cash cow without attempting to modernize(and without regard for further erosion of sales) or to build new facilities to replace the older plantswhile aggressively seeking to recapture the market share.

In the latter case, the Marketing Department forecasts that an additional 5 million pounds (about3.6 billion units) of penicillin G potassium (potassium salt of benzyl penicillin acid) will be

CD-A-II-39required by the year 2000; with 2 MM pounds by 1992; 3 MM by 1994; 4.1 MM by 1996; and 4.8MM by 1998. Penicillin G potassium presently sells for about $18 per pound ($25 per billion units).

If a new plant is to be built, the design will be based on state-of-the-art technology using highlymutated strains of Penicillium chrysogenum growing on glucose (corn syrup). A conventional batchprocess will be used unless adaptation to a fed-batch or continuous process appears feasible. Acrystalline product will be obtained after solvent extraction of the beer with amyl acetate or butylacetate.

As Director of Engineering, you have been asked to design the plant, determine the investmentrequired and assess the expected financial performance. You have also been asked to determine thecost-of-sales for the old plant at which it would no longer be competitive in profitability with a newplant.

A-II.5.3 Novobiocin Manufacture

(Robert M. Busche, Bio-en-gene-er Associates, January 1986)

Novobiocin is a general antibiotic produced by an aerobic fermentation of glucose by the organism

Streptomyces niveus. The basic elements of the process appear to be the fermentation of S. niveusin an appropriate medium of substrate and minerals, the adsorption of Novobiocin (as well as othernon-effective components expressed by the organism) on an ion exchange resin, and the desorption,concentration, and crystallization of a crude Novobiocin product consisting of 45% Novobiocin,21% Isonovobiocin, and 34% other similar molecules.

The Research Director of your large pharmaceutical company is interested in initiating research onproducing this product, but before committing funds, has asked you to evaluate the technoeconomicposition the company might develop in this new business. From very preliminary studies, itappears that the amount of Novobiocin made per fermenter batch is small, and that muchprocessing will have to be devoted to increasing yield and improving recovery efficiency. Also, asa result of low product concentration, oxygen transfer in the fermenter and power requirementsappear critical to the design and cost.

Your Information Specialist has developed the following literature references to serve as the basisfor your evaluation.

A-II.6.1 Polyvinyl Acetate Production for Polyvinyl Alcohol Plant

(Frank Petrocelli and Steve Webb, Air Products and Chemicals, January 2000)

A grass roots facility to produce polyvinyl alcohol (PVOH) is being constructed in a chemicalcomplex on the U.S. Gulf Coast. Your design team will complete the process engineering for theunit which produces polyvinyl acetate (PVAC). PVAC is further reacted in another part of thefacility to produce the PVOH final product. The polyvinyl acetate unit includes the polymerizationreactor system and the downstream recovery process. Your design must be capable of an annualproduction rate of 100 MMlb of PVAC intermediate.

PVAC is produced by the free-radical polymerization of vinyl acetate. Your company, PolyPenn,Inc., has experience and process knowledge using a continuous solution polymerization in whichthe solvent is methanol. The process uses a thermal initiator, which costs $5.00/lb. Thedecomposition kinetics for the initiator are given by the following expression:

To a first approximation, the polymerization follows classical free-radical polymerization kinetics

(as described by Flory; see references by Billmeyer and by Finch). Reaction conditions must bechosen to produce a medium-molecular-weight grade of PVAC, defined as a grade having anumber-average molecular weight of 130,000 (i.e., the number-average degree of polymerization,

CD-A-II-41Xn = 1,500 repeat units per polymer chain). Again, the references by Billmeyer and by Finchcontain mathematical expressions for determining the polymer molecular weight as a function ofreaction conditions. The reaction temperature must be maintained between 145 and 180°F, and thereaction pressure must be < 15 psig (this combination of conditions has been shown to reduce thepossibility of a runaway reaction in the event of a process upset). Also, for safety concerns (to limitthe amount of reacting material), the maximum size of any single reactor will be 10,000 gal.

Several decisions must be made in the initial design to choose among options for the process.Typically, the reaction does not proceed to complete conversion. The molecular weight of the finalPVAC is influenced by the level of conversion (higher conversion lowers molecular weight) andthe concentration of methanol in the reactor (increasing methanol lowers molecular weight). Thepolymerization can occur in a series of polymerization reactors. Your design team must decide onthe type of reactor (i.e., CSTR, PFTR, recycle loop), the number of reactors, reactor size, and themethod of heat removal (cooling jacket, cooling coil, and/or overhead condenser). Increasing thereactor size and the number of reactors can allow higher conversions for a given molecular weight,which would reduce recovery cost for the monomer. Obviously, there is a trade-off between therecovery cost and increased capital cost. Additionally, increased reactor size may reduce initiatoruse and cost. Your objective should be to find a design which achieves a minimum total cost overthe entire plant life.

After the polymerization reactors, the unreacted monomer must be removed from the polymerstream. In your company’s existing polymerization units, the monomer is removed in adistillation/stripping column. Methanol vapor is fed to the bottom of the column and a mixture ofmethanol and vinyl acetate monomer is taken as an overhead product. The PVAC exits from thebottom of the column in a methanol solution. To minimize product color formation, columntemperatures should not exceed 240°F. The bottoms from the PVAC/methanol column must have asolids content of 36 to 40% to be suitable for existing processing equipment downstream.

Some of the overhead product can be recycled and mixed with the reactor feed; the fraction whichcan be recycled is dependent on its composition. Excess overhead product is separated into purevinyl acetate and methanol in a separate, existing recovery process - assume a processing cost of$0.005/lb of recycle for this operation.

Ample cooling water is available at a supply temperature of 90 °F and must be returned no higherthan 110°F. Cooling water cost is $0.50/1,000 gal. Saturated steam is available at 150 and 600psig. The cost of steam is $5.00/MMBtu. Electricity is available at a cost of $0.05/kWh. Use themarket price for the cost of raw materials.

Finch, C. A., Polyvinyl Alcohol Developments, Wiley, New York, 1992.

Merk, W., et al., J. Phys. Chem., 84, 1694 (1980)

Nakajima, A., et al., J. Poly. Sci., 35, 489 (1959)

A-II.6.2 Butadiene to Styrene

(Bruce Vrana, DuPont, January 1997)

Butadiene (BD) is produced by the expensive extraction of BD from a crude C4 stream in an

ethylene plant. The BD value is about $0.06/lb when it is contained in the crude C4 stream, butabout $0.18/lb after it is extracted. Because of this price difference, processes are always beingsought to use the BD in the crude C4 stream without extracting it, and returning the remaining C4stream to the ethylene plant. A typical crude C4 stream has the following composition in weightpercent:

Dow has developed a process to dimerize the BD in a crude C4 stream to vinylcyclohexene (VCH)using a proprietary copper-loaded zeolite catalyst. The second step converts VCH to styrene viaoxidative dehydrogenation using another proprietary tin/antimony oxide catalyst.

Develop a plant design for a world-scale 1 MMM lb/yr styrene process using the new Dowtechnology, and determine the overall economics.

The plant design should be as environmentally friendly as possible. Recover and recycle processmaterials to the maximum economic extent. Also, energy consumption should be minimized, to theextent economically justified. The plant design must also be safe to operate (e.g., no flammable orexplosive mixtures).

Assume a U.S. Gulf Coast location. The BD contained in the crude C4 stream is valued at $0.06/lbin 1997 dollars, and any remaining C4s may be returned to the ethylene plant at no cost. Styrenesells for $0.30/lb. Oxygen may be purchased across the fence for $0.02/lb.

Patent Watch, Chemtech, 20 (May 1995).

U.S. Patent 5,329,057, July 12, 1994.

Because of the capacity limitations of urban landfills, biodegradable plastic packaging materials areof interest as a means to reduce the load on solid waste disposal systems.

Your research department has developed a mutant form of the bacterium Alcaligenes eutropus thatexpresses biodegradable poly (hydroxybutyrate) homopolymers and poly (hydroxybutyrate-valerate) copolymers. Although the copolymer has a lower melting point, it processes more easilythan the homopolymer. As a result, both may have value in plastic packaging. Under optimumconditions both the homopolymer and copolymer are produced at volumetric productivities ofabout 1.0 g/L-hr. Both products are best produced under phosphate limitation. The copolymer isproduced by adding n-propanol to the ethanol feed. The current research has been based on a fed-batch fermentation system. However, it has been proposed to use two-stage continuous culture inwhich the cells are first grown under conditions for optimum cell growth, followed by a secondstage under conditions optimum for product accumulation. Your research department is eager tomove ahead with the design of a commercial facility and will provide copies of appropriatereferences.

In the meantime, however, Dr. Douglas Dennis, an associate professor in the Biology Departmentof James Madison University, has cloned into a recombinant E. coli bacterium the genes thatcatalyze PHB formation in Alcaligenes. It appears that the new system produces polymer at therate of 2.7 g/L-hr. He has offered to provide an exclusive license to Imperial Chemical Industries(ICI) and will consult on a plant based on the recombinant organism.

As head of the ICI corporate plans department, you have been asked to evaluate the commercialpotential for developing a process to produce both homopolymers and copolymers at your plant atAtlas Point, south of Wilmington, Delaware. It is of interest to evaluate the economics of both ahomopolymer and a copolymer product and to suggest an optimum split, if one exists. Yourmarketing department has suggested a combined capacity of 50 million pounds per year for the firstplant. Either of the alternative processes could be considered.

A-II.6.4 Xantham Biopolymer

About 460 billion barrels of crude oil have been discovered in the United States to date, but only120 billion barrels have been recovered by primary gas drives or secondary water floods. A large

CD-A-II-44proportion of the remainder could be recovered, albeit at higher cost, by tertiary methods (enhancedoil recovery).

One such method involves the use of water-soluble polymers such as polyacrylamide to increasethe relative viscosity of sweep water to that of the crude oil so as to promote the mobility of theresidual oil in the reservoir. Polyacrylamide, although relatively cheap, does not possess the usefulproperties of polysaccharides such as xanthan gums, scleroglucan, dextran, etc. The biopolymersare injected at a rate of 1.4 to 1.7 lb/barrel of oil recovered. Excluding the polymer, the cost of thepolymer/surfactant flood amounts to $30 nd $40/barrel, including capital charges.

Your company, a major oil producer, is concerned about the rapid decline in productivity of itsCanyon Reef Reservoir in Kent County, Texas. Your Oil Production Department, which holdssome patents on producing xanthan biopolymers, is considering forming a joint venture with a foodcompany for developing and operating a fermentation facility to produce the 20 million annualpounds of polymer needed captively for a polymer flood of Canyon Reef. Merchant sales ofxanthan for food uses by the partner would also be considered if economically desirable.

Your Research Department has confirmed that xanthan can be produced from glucose by theorganism Xanthomonas campestris. Process and product data are summarized in the reports listedbelow.

Management has asked you to determine whether xanthan might be produced at a sufficiently lowprice to make the proposed EOR operation competitive with the importation of foreign crudes overthe next decade. Your Senior Vice President has also asked whether selling xanthan for currentfood uses would help to launch the new business at an earlier date than that compatible with EORmarket economics.

A-II.6.5 Rapamycin-Coated Stents for Johnson & Johnson

(Scott L. Diamond, University of Pennsylvania, January 2002)

In the treatment of heart disease, a common procedure involves balloon angioplasty to expand anarrowed coronary artery followed by placement of a metal support called a stent to keep the vesselopen. Stenting helps reduce vessel closure, a process called restenosis. However, even stentedvessels can undergo restenosis. There were 926,000 angioplasties in the U.S. in 1998 and 800,000

CD-A-II-46angioplasties outside the U.S. in 1999. Johnson & Johnson recently finished a clinical trial withpolymer-coated stents that slowly release the drug rapamycin. In 238 patients in Europe, not asingle patient had restenosis after 6 months with the rapamycin-coated stents. Johnson & Johnsonis positioned to obtain over 50% market share in the highly competitive stent market.

Production Criteria

1) Produce and purify medical grade Sirolimus (rapamycin) via batch bioprocessing using streptomyces fermentation. Determine how much rapamycin you must produce annually and how many batches will be necessary.

2) You will be provided with the metal stents from the Stent Manufacturing Group. You will carry out the drug-polymer coating of the stents and deliver the drug-polymer coated stents to the Catheter Manufacturing Group on a monthly basis.

3) You will buy pure medical-grade speciality chemical components for the polymer coating, but must develop the coating technology to achieve the correct drug loading and release characteristics needed in the clinical application. You will have to design a spray-coating process using ultrasonic nozzles as well as a drying process to remove the solvent. Solvent recovery is also required. Degradable polymers will include ε-caprolactone-co-glycolic acid.

4) Manufacture: 500,000 drug-polymer coated stents in year 1

1,500,000 drug-polymer coated stents in year 2 and after.

5) Estimate the capital cost and annual operating cost of the drug manufacture and coating systems.

References

www.uspto.gov U. S. Patent 6,153,252

U. S. Patent 6,273,913

www.ncbi.nlm.nih.gov/ Search pubmed: rapamycin stent

rapamycin streptomyces

Marx, S. O., and A. R. Marks, “Bench to bedside: the development of rapamycin and its applicationto stent restenosis,” Circulation. 2001, Aug 21;104(8):852-5. No abstract available.

A-II.7.1 R134a Refrigerant

(John Wismer, Atochem North America, January 2001)

A major shift is occurring in the fluorochemicals industry, particularly in that part of the industrywhich manufactures refrigerants. This involves the shift away from chlorine containing CFC’s(chlorofluorocarbons) and HCFC’s (Hydrochlorofluorocarbons) to HFC’s (Hydrofluorocarbons).This is because molecules containing chlorine degrade the protective ozone layer of the upperatmosphere. In automotive refrigerants, the shift has been away from R12 (Dichloro-difluoromethane) and towards R134a (1,1,1,2-tetrafluoroethane). This market is still growing asolder air conditioning systems are phased out around the world. Refrigerants use a nomenclaturewhich is universally accepted in the industry. A simplistic version involves the “rule of 90", inwhich 90 is added to the refrigerant’s numeric code. In the resulting number, the last digit denotesthe number of fluorine atoms, the second to the last, the number of hydrogen atoms, and the thirdfrom the last, the number of carbon atoms. When another digit occurs, it denotes the number ofchlorine atoms. When the compound is unsaturated, an extra digit is added to the left to indicatethe degree of saturation; “1" indicates a double bond in the molecule. The suffix letters denote theisomers based on symmetry considerations.

A major focus of the fluorochemicals industry has been to make use of retired HCFC or CFCmanufacturing equipment in the manufacture of new refrigerants. This project involves PennRefrigerants, a company with a fluorochemicals complex, which has several pieces of unusedequipment, particularly for distillation. It has a significant infrastructure for handling emissions,including an aqueous acid neutralization system, an incinerator for liquid organic wastes containingacids, and a thermal oxidizer for combustion of gaseous wastes. In other words, small wastestreams should not be a problem. It also has significant utilities infrastructure, including lowtemperature refrigeration (30 tons @ -40°C), a boiler plant capable of producing 150 psig steamwith 20K lb/hr of unused capacity, an electrical substation which can supply both 460V and 220V3-phase power, and a large excess of cooling tower capacity.

Penn Refrigerants is aware that there are several technologies available to manufacture R134a.They are considering licensing ICI’s patented process. You (Quaker Consultants) have beenapproached to evaluate the capital required to retrofit the Penn Refrigerants plant on the Gulf Coastto make R134a using the ICI technology.

The ICI process is documented in U.S. Patent 5,382,722. It involves two reaction steps:

TCE + 3HF → R133a + 2HCl (1)

R133a + HF → R134a + HCl (2)

Not mentioned in the patent, but implied, is that gas phase reaction (2) has a relatively severeequilibrium limitation. Its heat of reaction is about 6.5 kcal/mol (i.e., endothermic) and the entropyof reaction is about -2.5 cal/mol-K. Also, the patent mentions a R1122 impurity which boils in the

CD-A-II-48same range as R134a. This is the most troublesome olefin, but there may be others. One way todestroy these olefins is with chlorination technology. Penn Refrigerants has chlorine storage andfeed systems available in their plant. Chlorination can be accomplished photochemically orperhaps, more simply, catalytically. The R134a molecule is resistant to chlorination at thetemperatures used to saturate the double bond. The saturated chlorine-containing compound ismuch less volatile than R134a.

Penn Refrigerants has placed constraints on its plant:

Gaseous HF or HCl cannot be compressed.

HCl must be recovered by distillation and absorbed

into aqueous form at 36% concentration.

Inconel 600 or better is required for reactor and HF

reboiler service

There are useful VLE data for mixtures of HF, R133a, and R134a in the Journal of FluorineChemistry, 61, 123-131 (1993). Some LLE data are in European Patent No.0 509 449 A2. Hydrogen fluoride has some odd thermodynamic characteristics which can makeequipment design of HF systems tricky. A good guess at its enthalpy chart with a good discussionappears in a paper by Yarboff and Lightcap (J. Chem. Eng. Data, 9, 2, 178, 1964). ASPEN PLUSuses a special equation of state to approximate the HF association effects. Does this approximationagree with the Yarboff and Lightcap chart? If not, how might this affect your design?

U.S. Patent 5,382,722.

Journal of Fluorine Chemistry, 61, 123-131 (1993).

European Patent No. 0 509 449 A2.

Yarboff and Lightcap, J. Chem. Eng. Data, 9, 2, 178 (1964).

A-II.7.2 Biocatalytic Desulfurization of Diesel Oil

(Robert M. Busche, Bio-en-gene-er Associates, January 1994)

The EPA's revised pollution guidelines for on-highway diesel fuels took effect on October 1, 1993,and additional Clean Air Act amendments are pending. As a result, the sulfur content of diesel fuelwill have to be reduced from 1 to 2% down to 0.05% as compared with 0.3% conventionallyattainable with high-pressure hydrodesulfurization.

For a number of years researchers have attempted with little success to develop a biological systemto remove organic sulfur. However, in 1989, J.J. Kilbane at the Institute of Gas Technologysucceeded in isolating a bacterium that oxidized dibenzothiophene to 2-hydroxybiphenyl andliberated sulfur.

Based on this discovery, scientists at Energy Biosystems Corp. of Houston, TX, have beendeveloping a biocatalytic desulfurization process using the bacterial enzyme IGTS8 to catalyze thereaction in a CSTR bioreactor. The extracellular enzyme is produced by the bacterium in an aerobicfermenter. The enzyme is then transferred as a supernatant solution to the bioreactor, where itmixes with high-sulfur diesel oil, oxygen, and other process chemicals. In the reactor, the sulfur isenzymatically removed from the oil to levels substantially below the 0.05% new regulatory limitwithout reducing fuel value. After reaction is completed, the water/oil emulsion is fed to a separatorand the desulfurized oil is transferred to storage. The aqueous phase is sent to a separations unit toremove the sulfur, after which the enzyme/water mixture is recycled to the bioreactor after taking apurge of spent enzyme.

Dr. Daniel J. Monticello, VP Research for EBC, has recently invited your oil company to join aconsortium to develop the process to the point of commercialization. Before deciding to accept thisinvitation, the director of your Refining Division has asked you to evaluate the expected economicsof the proposed process as compared with the demonstrated costs for hydrodesulfurization in the30,000-BPD diesel unit in your Richmond, CA, refinery. You are asked to identify the major cost

CD-A-II-50elements and assess the sensitivity of cost to process improvements that might be effected withfurther research on economically critical process parameters.

A-II.7.3 Sulfur Recovery Using Oxygen-Enriched Air

(Mark R. Pillarella and Rakesh Agrawal, Air Products and Chemicals, January 1993)

The Clean Air Act, passed in 1990 and scheduled to become effective in 1995, will force chemicalcompanies to reduce their emissions, into the atmosphere, of many environmentally detrimentalchemicals. These include sulfur which occurs as H2S in sour natural gas and refinery gas.Recognizing that sulfur recovery is a fast-growing business, you have recently formed your ownengineering company, SULFREC, which specializes in sulfur recovery. A small chemicalcompany has requested that SULFREC submit a bid to design a process for removal of sulfur froma 23-metric-ton-per-day gas stream (90 wt% H2S and 10 wt% CO2) using the modified Clausprocess described below. The principal reactions are:

The H2S - CO2 gas is at 38°C and 1.72 bara. Ninety-five percent of the H2S is converted to sulfur.The sulfur recovery system is to be installed in Houston, Texas.

The modified-Claus process typically uses air as its oxygen source. However, O2 -enriched air mayprovide a more economical alternative. Your company has decided to investigate three alternativedesigns, each using the modified-Claus process, but with different oxygen sources:

1. Ambient air

2. O2-enriched air using a membrane

3. O2-enriched air using vacuum swing adsorption (VSA)

CD-A-II-51The company requesting the bid has stipulated that a comparison of the three alternatives, as wellas a full design of the most economical process, be submitted.

Information on the modified-Claus process is available in the literature. Information for the designof the membrane and VSA processes will be supplied by Mark Pillarella at your request.

CD-A-II-52A-II.7.4 California Smog Control (E. Robert Becker, Environex, January 1995)

Background

A primary gaseous air pollutant from combustion sources such as power plants is oxides ofnitrogen (NOx). Since NOx is a known precursor to ozone formation, the Clean Air ActAmendments of 1990 call for reduction of NOx from certain facilities throughout the United States.

You are a project engineer for an independent power producer in California. The state hasmandated that your company reduce NOx emissions from your Los Angeles facility by 200 tons peryear.

The LA facility currently operates two units, a 25-MW combustion turbine and a set of four 3-MW(12 MW total) diesel engines. The diesel engines share a common exhaust stream. The NOxreduction can come from either unit or both.

You are to design the NOx removal system for each unit and determine which of the two systems isthe most cost effective for NOx removal.

The primary form of NOx for a combustion source is nitric oxide (NO). The NOx removal systemto be considered is Selective Catalytic Reduction (SCR).

SCR System Background

SCR removes NOx by reacting it with gaseous ammonia (NH3) at about 700°F in the presence of acatalyst according to the reaction:

4 NO + 4 NH 3 + O2 → 4 N2 + 6 H2 O

Kinetics

The rate of reaction is first order in NOx and the overall apparent rate constant (combination ofmass transfer and reaction rate) is 43,000 1/hr at 700°F.

A typical SCR system has four major components:

• A liquid ammonia storage tank

• An ammonia vaporizer• An ammonia injection grid (to evenly disperse the NH3 across the duct)• A catalyst

The ammonia injection grid must be designed to evenly distribute the ammonia across the ductwithout restricting flow. Since NO concentrations vary across the duct, the ammonia injection gridmust also be designed such that ammonia injection can be adjusted to match the NO concentrationsacross the duct of the exhaust system.

CD-A-II-53Several types of catalysts are used for SCR. The catalyst for this design is an extruded, squarepitched homogeneous catalyst. A schematic of a catalyst brick is given in the figure.

pitch

b 6 in. a

6 in.

Figure SCR Catalyst Brick

The catalyst bricks are 6 by 6 inches and can be cut up to 3 feet in length. Catalyst bricks are thenarranged side by side and front to back as necessary to achieve the desired dimensions.

Part of your assignment is to design the ammonia injection grid, catalyst and catalyst housing.Specify the number and arrangement of catalyst bricks and total catalyst volume. The ammonia

CD-A-II-54injection grid is composed of a series of pipes with holes or nozzles to inject the NH3. Keep inmind that the grid should be designed to evenly distribute the NH3 across the duct and be flexibleenough to adjust to match NO concentration variations.

Compare the cost of the two units to determine which is the economic choice. Compare bothcapital costs and NH3 consumption over a 10-year catalyst life and the incremental cost on thepower generated.

For many years your plant on the Texas Gulf Coast has produced tetrahydrofuran (THF) for use asa synthetic fiber intermediate. The reaction is carried out in water solution, producing a crude THFwhich also contains lower aliphatic alcohols as byproducts plus some gamma-butyrolactone (GBL),which is an unreacted intermediate. The THF is purified in a three-column distillation train. Theimpurities have been incinerated or sent with the water to the biological effluent treatment system.

Last week (without consulting the technical staff) your company's Board of Directors issueda press release stating that the plant is to be converted to a "Zero Emissions" operation by January1, 1994. Your boss, the Chief Engineer, practically had a coronary on the spot, but he recovered intime to assign the job to your team. He also scheduled you to make a presentation to the Board ofDirectors on April 30, to outline your recommendations and present the economics of the variouspossible solutions. In practical terms your job is to reduce emissions to the lowest possible level,but to do it in the most cost-effective manner.

Some of the ground rules are:

• The three existing columns are not to be modified.

• Fluegas is considered to be an emission (greenhouse effect), so incineration is not

acceptable.

• The cooling tower is not included in the Zero Emissions envelope.

• The biological treatment system is not an acceptable solution for the waste water.

If you can't burn it and can't discharge it, what can you do with it?

A local solvent supplier has offered to buy any of the alcohols which meet the purityspecifications shown below. He has quoted the following prices for tank truck quantities:

99.8% Methanol 11¢/lb 99.5% Propanol 29¢/lb

99.7% Ethanol 29¢/lb 99.9% Butanol 54¢/lb

As an alternative, any mixture of alcohols can be sold as a gasoline additive for 9¢/lb, as long as itcontains no more than 0.5% (wt) water.

THF can be recycled to the crude THF tank at an operating cost savings of 25¢/lb, and GBLcan be recycled to an earlier step in the process at a savings of 15¢/lb. Any water and impuritieswhich accompany these recycles must be reprocessed through the distillation train, but you mayassume the existing columns can handle this.

Waste water can be used as cooling tower makeup, as long as its organic content is below50 ppmw. It will replace raw water at a price of 35¢/Mgal.

The following utilities are available for your use:

150 psig steam, dry & saturated, @ $3.25 per Mlb.

A-II.7.6 Volatile Organic Compound Abatement

(E. Robert Becker, Environex, January 1994)

The 1990 Clean Air Act requires the reduction of volatile organic compound (VOC) emissions. AllVOC emission sources of 10 tons/year or greater are required to retrofit abatement processes usingthe best available control technology (BACT).

A paint spraying plant emits VOCs from the vent of its paint spray booths. The stream containsprimarily toluene, methyl ethyl ketone (MEK), and xylene, with small impurities of silicone andphosphorus. The concentration of VOCs in the dryer effluent varies between a minimum of 0.3wt% VOC and a maximum of 1.2 wt% VOC with an approximate composition of 50% toluene,25% MEK, and 25% xylene.

You are commissioned by the painting company to evaluate three alternative technologies for VOCreduction: thermal incineration, catalytic incineration, and carbon adsorption of the VOCs followedby destruction. A nearby bottle washing plant can use low-quality steam.

Design an emission control plant for 50,000 scfm of vent gas at 100°F and 25% relative humidityfor 99% removal. The plant is located in Dearborn, Michigan, and the paint spray booths operateon a single 12-hour shift per day. Include the necessary start-up controls. The available fuel isnatural gas or oil. Calculate the capital and operating cost and the $/lb or ton of VOC removed.Compare the three processes and recommend which is most suitable for this application.

Your company, BIG-D CHEMICALS, is a major producer of pentafluoroethane (CF3CHF2), which

is also known as hydrofluorocarbon 125 or HFC-125. HFC-125 is one of the new ozone-friendlyfluorocarbons, and it is a replacement for chloropentafluoroethane (CF3-CClF2) or CFC-115 inmany refrigerant applications.

In the production of HFC-125, some CFC-115 is produced, and this material must be removed fromthe HFC-125 product. In addition, hydrochloric acid (HCl) is always produced as a byproduct, andit must be recovered as a reasonably pure stream for the process to be attractive.

Your new job with BIG-D is to find the most economical process to recover HFC-125 from amixture which contains HFC-125, HCl, and CFC-115. The HFC-125 product must contain nomore than 100 ppm-wt of other organic impurities (e.g., CFC-115, HCFC-124, etc.) and the aciditylevel (as HCl) must not exceed 10 ppm-wt. In addition, the process will be more economicallyattractive if you can recover anhydrous HCl which contains no more than 10 ppm-wt of organicimpurities. If you are unable to meet the anhydrous HCl purity specification, the HCl must beabsorbed in water (35 wt%) and subsequently air stripped to remove the organic impurities.Aqueous HCl solutions are a drug on the market and have essentially no value; the absorption routeis used only to avoid neutralization and waste disposal costs. Organics in the air stripper offgasmust be collected and disposed of in an environmentally acceptable manner (e.g., incineration).

For the process to be economical, CFC-115 must be recovered and recycled to the reactor.Although recycle CFC-115 may contain up to 5 wt% HFC-125, there is a cost penalty associatedwith HFC-125 recycle, so you will probably want to minimize HFC-125 in the CFC-115 stream.The recycle CFC-115 may also contain up to 1 wt% HCl; there is no cost penalty associated withHCl recycle at this level.

The composition of the feed stream to the new recovery/purification process is: HFC-125 5,000pph, CFC-115 500 pph, HCl 2,000 pph (available as a saturated vapor at 275 psig). The feedstockvalue of this stream is $2.50/lb.

The values of the various product and byproduct streams are as follows:

BIG-D's fluorochemicals facility is located on the U.S. Gulf Coast. The new plant will be situatedadjacent to an existing fluorochemicals manufacturing plant and will share some common facilities(i.e., control room, maintenance shops, technical office building, etc.). Storage facilities exist forboth anhydrous and aqueous HCl. Except for the above, all equipment will be new (i.e., there is noused/existing equipment available for your use). You can assume an operating utility of 85% (7,446hours per year) for both new and existing facilities.

CFC-115 (nbp = -39.1°C) and HFC-125 (nbp = -48.1°C) can be removed from HCl (nbp = -77.5°C) by conventional distillation; this process is energy intensive and requires low temperatures,but it has been demonstrated in the laboratory. BIG-D's research people have been very creativeand have also developed an extractive distillation process for recovering HFC-125 and makinghigh-purity anhydrous HCl. The extractive distillation process requires more equipment but usesless energy. Potential extractants are HCFC-123 (CF3-CHCl2), which is valued at $3.00/lb, andHCFC-124 (CF3-CHClF), which is valued at $3.50/lb. These materials are available on site aspressurized liquids at 10°C and 100 psig.

Your assignment is to develop both conventional and extractive distillation processes forrecovering HFC-125 and HCl from the specified feed mixture. You will need to develop optimumflow sheets, size and cost equipment for each case, and compare the economics of the twoprocesses. Your flow sheets should include energy recovery (heat integration) as appropriate. Youwill also need to develop a control strategy for your preferred case; the control scheme shouldaddress start-up and shut-down conditions as well as steady-state operation.

Notes

1. CFC-115, HFC-125, and HCFCs 123 and 124 are nonflammable and noncorrosive. Carbon steel is a satisfactory material of construction for pressure vessels; if the temperature is less than 0°C (either operating or upset conditions), a Charpy impact test is required. HCl may be handled in either stainless steel or low-temperature carbon steel (Charpy impact tested) equipment.

2. On the U.S. Gulf Coast, cooling towers will supply water at about 31°C in the summertime. This should be the design basis for any water-cooled condensers or heat exchangers. The CTW

CD-A-II-59 supply temperature is about 10°C during the coldest months. CTW is high in chlorides (due to evaporation) and is quite corrosive.

3. The largest distillation column on the plant site is 150 ft tall. It was designed by the Plant Manager when he was a junior engineer a number of years ago. He is quite proud of this column, and he often points it out to new visitors to the site. You probably don't want to change this (or his feelings about you).

4. If any of the new process steps operate under vacuum, you should assume there will be air leakage into the process. While this is not a safety hazard, you will need to include facilities to remove inerts from the HFC-125 product.

5. Purity requirements for the new HFC products are much more stringent than for your current CFC products. As a result, analytical techniques have not yet been fully developed to analyze for low levels of some trace impurities. BIG-D's analytical chemists are currently working to develop more sensitive analytical methods to identify other impurities.

6. Thermodynamics/physical property information will be provided for the chemical species

Although reducing the concentration of carbon dioxide in the atmosphere and its concomitantgreenhouse effect has become an increasingly important public issue, little progress has been madebecause the demand for electric power based on fossil fuels continues to grow. Worldwide, one-third of all carbon dioxide emissions come from electric generating plants. Emission levels can belowered in one of three ways: (1) converting to alternative non-fossil fuels such as nuclear orbiomass; (2) increasing the energy efficiency of the fossil fuel-based process; or (3) preventingcarbon dioxide in the flue gas from reaching the atmosphere.

In Japan, Mitsubishi Heavy Industries, Ltd., and Tohoku Electric Company, Inc., have beenexperimenting with the use of microalgae to fix carbon dioxide in stack gas for subsequent recycleas a solid fuel. In this process, the algae, Nannochloropsis salina and Phaeodactylum tricornutum,are grown in sea water contained in shallow lagoons under an atmosphere of flue gas containing 10to 12% carbon dioxide. The nutrients NaNO3 for nitrogen and NaH2PO4 for phosphorus are addedin small concentrations. After harvesting, the microalgae is dried and recycled to the power houseas a solid fuel. The kinetics of the process were defined in the study.

As Branch Chief for the Department of Energy's Office of Carbon Dioxide Emission Control, youhave been asked by the Deputy Secretary to evaluate the possible use of this approach in designing

CD-A-II-60the emission control facilities for the proposed 600-megawatt generating station to be built betweenLos Angeles and San Diego to service the expanding needs of these communities. Governmentland can be made available for this purpose.

Specifically, you are asked to evaluate the cost and investment for an algae facility compared withthe best alternative of your choice for reducing emissions by 50%. Alternative approaches toemission reduction are evaluated in the first four references. Your comments on the efficacy ofother alternatives will also be of interest. Results should be expressed in terms of $/kWhr ofgenerated electricity. Please test the sensitivity of cost to the levels of emission reduction for theapproaches you consider.

Likewise, since the most effective way to reduce emissions is to increase the energy efficiency ofthe generation process, please ascertain the potential equivalent cost reduction vis-a-visimprovements in energy efficiency, and hence, determine the limiting minimum power cost at100% efficiency. If data are available for the generation station under evaluation, determine thepertinent potential cost savings due to improvements in the energy efficiency.

A-II.7.9 Hydrogen Generation for Reformulated Gasoline

(E. Robert Becker, Environex, January 1994)As a result of the Clean Air Act Amendments, hydrogen consumption within refineries willincrease and hydrogen byproduct production from catalytic gasoline reforming will decrease. Thisincreased use will be brought about by the required reduction of benzene, olefin, aromatics, andsulfur in gasoline and the reduction of aromatics and sulfur in diesel fuel. This demand will, inmost cases, be met by on-site facilities for the production of hydrogen.Hydrogen is currently produced by either steam reforming of methane or by partial oxidation ofmethane with high-purity oxygen and steam. Your research department has developed a new

CD-A-II-61autocatalytic reactor using air, methane, and steam that has some very definite advantages. Theyare: 1. Refinery use will require high-purity hydrogen (99.9%) to minimize inert build-up in the recycle hydrogenation processes. In the autocatalytic process, the hydrogen separation is much easier than the nitrogen-oxygen separation and is less energy demanding than the reforming operation with steam. 2. The catalytic process operates at a lower temperature than is required for the steam reforming (1,000˚C as compared with 1,400˚C for the Shell/Texaco process). Moreover, the equipment is much simpler. The process can also be operated at higher pressures, thus saving on compression of the product hydrogen.In the proposed process, methane, steam, and air are each preheated to 600-700˚C and fed into acatalytic reactor containing a bed of refractory nickel catalyst. Initial combustion results in atemperature up to 1,000˚C. The gas passes through heat exchange and a heat recovery boiler beforeentering a multi-stage CO shift converter. The gas then passes through a CO2 wash tower to acryogenic separation unit where the hydrogen is separated from the nitrogen, argon and methane.You are required to prepare a preliminary cost estimate for a plant to produce 50 MM scfd 99.0%hydrogen at 30 atm. The following design data should be used for this evaluation.H2O/CH4 feed ratio > 2.0 to prevent carbon formationCO shift conversion is 99% of equilibrium.Cost Data

The refractory nickel catalyst is a spherical pellet of 0.005 m diameter. The catalyst bed has a voidfraction of 0.48 and a bulk density of 1200 kg/m3. The catalyst must be replaced annually.Kinetic Data

The treated effluent stream must satisfy the following limits:

COD = 300 mg/L F = 5 mg/L

where COD is chemical oxygen demand, TOC is total oxygen content, and TSS is total suspendedsolids.

There are no restrictions on the method(s) you select for remediation (e.g., physical, chemical,biochemical, leading, etc.). However, it is desirable not to generate much additional waste in theclean-up process. It is also desirable to find modifications that reduce the waste generation tolevels that satisfy the effluent limits. Note that your company has sufficient capacity to store the

CD-A-II-63contaminated stream for one month. Your design report should address health and regulatoryissues.

The manufacturing process for making optical fibers involves high temperature oxidation of silicontetrachloride (SiCl4) to form glass particles (SiO2 and GeO2) that are incorporated into a glasspreform rod. This rod is subsequently drawn in a furnace to produce optical fiber. Germaniumtetrachloride is added to increase the refractive index of the glass core in the optical fiber preform.It is known from experimental studies that the oxidation of GeCl4 to GeO2 proceeds to only 25%completion whereas oxidation of SiCl4 is nearly complete. In addition, particle deposition is only50% efficient, resulting in further losses of germanium. Due to this loss and the high cost ofgermanium, a need exists for developing a process to recover germanium from optical fibermanufacturing effluents. For environmental reasons, the process design must also provide for theremoval of chlorine and particles.Your company currently operates with 50 preform manufacturing units. Each unit is equipped witha small packed-column scrubber that is known to be underdesigned based on the current effluentproduction rates. The scrubbing solution is not recirculated and there is no recovery of germanium.Your engineering group has been designated to prepare a process design for a new scrubbingsystem to efficiently remove GeCl4, Cl2, and particles from the effluent stream:

GeCl4 200 g/min

SiO2 75 GeO2 1 Cl2 375 O2 7

The new scrubbing system should remove 99% of both GeCl4 and Cl2. You should also design asystem to recover germanium and convert it to GeCl4.In the existing process, vapors of SiCl4 and GeCl4 in an excess of oxygen are introduced into theoptical fiber preform production units where the following reactions occur at high temperature:

SiCl4 + O2 = SiO2 + 2Cl2 (1)

GeCl4 + O2 = GeO2 + 2Cl2 (2)

Both reactions reach equilibrium which corresponds to 100% completion for reaction 1 and 25%completion for reaction 2. Incorporation of solid particles into the glass preform rod is only 50%efficient. The effluent stream therefore contains SiO2 and GeO2 particles, unreacted GeCl4 and O2,and the reaction product Cl2.Currently, effluents from each preform production unit are drawn into small (0.25 m diameter, 0.5m high) packed bed scrubbers. The scrubbing liquid is an aqueous NaOH solution adjusted to pH13. A single fan unit draws the effluents into the scrubbers. Due to operating requirements, it isnot possible to make a tight seal between the effluent stream outlet and the inlet to the scrubbingsystem. Hence, the effluent stream gets diluted with a large amount of room air as it enters thescrubber.

CD-A-II-64Within the scrubbers, GeCl4 and Cl2 are removed from the gas stream by absorption and convertedto soluble species according to the following reactions:

GeCl4 + 5OH- = HGeO 3- + 4Cl- + 2H2O (3)

Cl2 + 2OH- = ClO- + Cl- + H2O (4)

The particles dissolve according to:

GeO2 + OH- = HGeO3- (5)

SiO2 + OH- = HSiO 3- (6)

Hydrogen peroxide (H2O2) is also added to the system to reduce the hypochlorite concentrationaccording to: ClO- + H2O2 = Cl- + H2O + O2 (7)The R&D department has found that germanate (HGeO 3- ) can be quantitatively removed fromsolution by precipitating with a divalent cation such as Mg2+ according to: Mg2+ + HGeO 3- = MgGeO 3-(ppt) + H+ (8)Similarly, silicate ions are precipitated according to: Mg2+ + HSiO 3- = MgSiO3(ppt) + H+ (9)Experiments have shown that HGeO 3- and HSiO 3- are precipitated equally well and that a moleratio of 1.25 to 1 for Mg to total of Ge + Si is required to precipitate 100% of the Ge and Si. Mgcan also be precipitated as Mg(OH)2 according to: Mg2+ + 2OH- = Mg(OH)2(ppt) (10)The R&D department has also measured the solubility of Ge and Si in scrubbing solutions atvarious pH values. The results are presented in a table (available from the AIChE) and may beuseful in the design of a germanium recovery system.MgGeO3 can be used as a feed to make GeCl4 which, after purification, can be used in opticalfiber production. The tetrachloride is formed according to: MgGeO3 + 6HCl = GeCl4 + MgCl2 + 3H2O (11)Additional reactions that may also take place include:

MgSiO3 + 6HCl = SiCl4 + MgCl2 + 3H2O (12)

A-II.8.3 Solvent Waste Recovery

(David G. R. Short, DuPont, January 1997)

Your company operates a polymer-processing facility which has three major waste streams. Whilethere have been no major problems with the regulatory agencies in the past, the new CEO wants allfacilities to have an environmentally friendly image. And there is a definite smell from your bio-pond which the neighbors complain about when the wind shifts. In addition, the existing permitsare coming up for renewal. While the negotiations are seen as friendly, the expected outcome is torenew the current permit provided that an improved waste system is in place by the year 2002. TheCEO thinks there is an economic incentive to have the facility running sooner.

2. Water 1 stream: 100 gpm Water with 2 percent acetaldehyde.

3. Water 2 stream: 100 gpm water with 3 percent ethylene glycol.

Current Facilities

1. The air stream is vented to the atmosphere.

2. The water wastes are sent to a bio-pond. The waste chemicals are oxidized to CO2 and H2O. The pond is sparged with air to provide oxygen for the bacteria. The pond is at capacity. The holdup time for the complete oxidation of the wastes is 36 hours.

Specifications

1. > 90% removal of all contaminants from the waste streams. This includes any new vent streams which may be created in the waste facility.

2. If solvents are recovered for reuse, the purity must be at all times > 99.5% pure with water as the major impurity.

3. The waste-handling system must never shut down the production plant.

Upset Conditions

1. The air flow can decrease by as much as 50% in 30 seconds. The total contaminant load will stay the same.

CD-A-II-66 2. The Water 1 stream has shown short-term flow rate fluctuations of 10% with no change in contaminant concentration.

3. The Water 2 stream has shown short term fluctuations of as much as 50% with a 2 × increase in the contaminant concentration.

Expansion Plans for the Polymer Facility

1. There is a high probability that the air stream will double in size, but the contaminants will increase by 50%.

2. The Water 1 stream will most likely have the same flow rate, but the concentration may be as high as 5% acetaldehyde.

3. The Water 2 stream will most likely double, but the concentration will be cut to 2% ethylene glycol.

Assignment

Design a facility that will meet the above specifications. Include in your study:

1. A discussion of alternatives to your final process.

2. A detailed development of your selected process. 3. A demonstration that the process is operable, using a model that shows the process can be started up, operated, and shut down. 4. A demonstration of the process sensitivity to changes in feed conditions.

A-II.9 ENVIRONMENTAL – SOIL TREATMENT

A-II.9.1 Phytoremediation of Lead-Contaminated Sites

(Robert M. Busche, Bio-en-gene-er Associates, January 1995)

A large chemical company quartered on the East Coast spends about $400 million annually toremediate contaminated aquifiers and sites associated with past manufacturing operations. Much ofthis is spent on sites contaminated with lead from the manufacture of tetraethyl lead, lead-basedpaints, and lead cartridges. For example, the soil of a 25-acre site within a large plant located inNew Jersey contains as much as 2,000 ppm of lead as inorganic salts to a depth of 2 feet. Thedistribution is as follows: at the surface, 2,000 ppm; 6”, 1,000 ppm; 12”, 500 ppm; 24”, 0 ppm.

A straightforward, albeit expensive way, to remediate this site would be to excavate the top 2 feetof soil and replace it with uncontaminated fill, then mix the contaminated material with cement tostabilize the lead, and dump it into a landfill. The cost of the remediation alone was estimated to

CD-A-II-67amount to about $9 million. An additional $1.8 million would also be required for documentation,sampling, analytical tests, decontamination, etc.

has been experimenting with hyperaccumulating plants that can absorb lead and other heavy metalsat up to 2% of the dry biomass weight. Such plants can produce 20 tons (dry basis) per acre perharvest of vegetative matter. Cunningham has identified as candidate crops a perennial plant, hempdogbane, which yields a single annual crop, and two annual plants, brassica (Indian mustard) andcommon ragweed. Brassica can be planted in the spring and fall to yield two crops annually.Ragweed is planted in early summer and harvested in the fall before blooming. Operators involvedin the planting, harvesting, and handling of the biomass are required to wear Level C personnelprotective equipment (PPE), e.g., hooded, unlined Tyvek coveralls, goggles, and masks.

After harvest the biomass can be treated in a number of ways:

1. It can be incinerated to reduce its volume by 75%. The ash is then stabilized with cement and landfilled; or

2. It can be chopped, dried, and then fed pneumatically to a flame reactor as developed by the Horsehead Resource Development Company of Monaca, Pennsylvania. Natural gas and oxygen-enriched air (40 to 70% oxygen) are fed to the reactor with the biomass to produce a very hot reducing gas at 2,000°C. Under reducing conditions the biomass is consumed to produce carbon monoxide while the lead compounds are reduced to elemental lead vapor. Small amounts of biomass minerals and dirt from the harvest operation are melted into slag, which flows down into a horizontal separator where it is separated from the lead vapor. The nonhazardous slag is tapped, cooled, and disposed of in a landfill. The lead vapor is passed into a combustion chamber, where it is converted into the oxide, cooled with cold air, collected in a bag house, and stored. (It is extremely difficult and expensive to quench lead vapor without converting to the oxide.) The crude oxide is sold to a lead smelter at about 50% of the price for pure lead (currently 38 to 40 cents per pound); or

3. It can be chopped and fed with suitable nutrients to an anaerobic digester wherein 95% of the carbohydrate is converted to a mixture of carbon monoxide and carbon dioxide. In the process the lead precipitates as lead sulfide and is centrifuged to separate it from the residual carbohydrate and water. The aqueous layer is recycled to the digester after taking a suitable purge to bleed off salts. The aqueous waste can be treated with lime to precipitate the salts. As with alternative 2, the lead sulfide can be sold to a lead smelter at 50% of the price for pure lead.

As a member of your corporate plans department, you have been asked to evaluate thetechnoeconomic position of phytoremediation and recommend an appropriate plan of action forremediating the site to the plant manager. Cunningham has promised to provide additionalinformation and, perhaps, a tour of his laboratory. John Pusateri of Horsehead will perhaps providea tour of his pilot facilities.

A-II.9.2 Soil Remediation and Reclamation

(1993 Environmental Design Contest Problem)

A large area (hundreds of square miles) in an arid region of the Pacific Northwest has beencontaminated with fallout from a neighboring manufacturing region. The site is to be bothremediated and reclaimed. Remediation will be defined as reducing the concentration of identifiedcontaminants below the threshold values listed below. Reclamation will include the use of a waterharvesting system to enhance the growth of natural vegetation or agricultural crops on the site.Water harvesting is an ancient concept that has been applied to increase biomass production in aridand semi-arid lands. Water harvesting concepts currently applied in arid lands continue to besomewhat primitive technically and small in scale. The desire to improve the technical state ofwater harvesting by finding, selecting, designing and/or testing sealants for catchment areas and bydeveloping equipment that can apply the sealants on very large areas (hundreds of miles) inrelatively short times (a few years) in a cost effective manner. These water harvesting catchments(sealed areas) may be tied into no drainage growing strips.The design team is given the following three tasks:1. Develop and demonstrate a bench scale process to remove the identified contaminants from a 5- kg sample of the soil. Develop a conceptual design of the process applicable to the field-scale project.2. Develop and demonstrate a non-geomembrane, UV-resistant, water-repellent, erosion-resistant, sealant that can be applied to the soil surface as an aid in water harvesting.3. Develop a conceptual design for a machine capable of applying the sealant developed in task 2 to the surface of the remediated area.

Regional and Soils Description

The area to be remediated is located adjacent to a large river in an arid climate. Natural vegetationis desert shrub and bunch grasses. Soil material is the result of catastrophic flooding followed by

CD-A-II-69deposition of river alluvium. The resulting material is a mixture of cobbles and sandy loam soilmaterial. The following particle size description is typical of the material at the remediation site.

SOIL

Component Size (mm) Weight Basis Weight Basis

(%) (%) < 2mm ______________ ______________ ______________ ______________ Large cobbles 305 15 None Small cobbles 152 30 None Fine pebbles 8.5 5 None Very course sand 2.000 10 20 Course sand 1.000 7 14 Medium sand 0.500 7 14 Fine sand 0.250 5 10 Very fine sand 0.106 4 8 Silt and clay 0.050 17 34 TOTAL 100 100To approximate this material with a soil from New Mexico, we have chosen the soil series Casito(Petrocalcic Ustollic Paleargrid). This soil is formed in alluvium at the base of mountainwatersheds. It is found on alluvial fans and terraces. It contains a similar mixture of cobbles andfines as the remediation site; however, the source of the alluvium is storm runoff rather thancatastrophic flooding and river deposition. Provided below is some additional information on thecontaminants in this soil.

A-II.10 ENVIRONMENTAL – MISCELLANEOUS

A-II.10.1 Fuel Processor for 5 KW PEM Fuel Cell Unit

(Jianguo Xu and Rakesh Agrawal, Air Products and Chemicals, January 2002)

Fuel cell technology is considered to be a disruptive energy technology. Fuel cells use fuel in anelectrochemical combustion process that converts the chemical potential of the fuel with respect tothe combustion product directly into electrical power. They are more efficient and moreenvironmentally friendly than conventional energy technologies. Fuel cells, especially the protonexchange membrane (PEM) fuel cell, are being considered for distributed power generation (DG).Using a fuel cell for DG reduces the energy loss due to power transmission, and can eliminatepower outages due to weather-related or other causes. It also allows for efficient use of the low-level waste heat from the power generation process. This low-level heat can be used for producinghot water, and for room heating. Since the PEM fuel cell uses hydrogen gas as fuel, a supply ofhydrogen gas has to be installed for a fuel-cell power generator to work.

Hydrogen for use in residential fuel cells can be produced from pipeline natural gas using a fuelprocessor. Assume that a residential, fuel-cell, electric-power generator with 5 kW electricityoutput has an efficiency of 50% (the electricity output from the fuel cell is 50% of the lower

CD-A-II-71heating value of the hydrogen consumed in the fuel cell). The desired hydrogen pressure is 0.5barg. Note that the CO content in the hydrogen supplied to the fuel cell must be below 10 ppm, andthe sulfur content must be less than 0.1 ppm. Nitrogen, carbon dioxide, methane, water vapor, andother inert gases are not poisonous to the fuel cell. For design purposes, a fuel gas with less than 3vol% of hydrogen cannot be used to fuel the fuel cell.

A possible approach: Natural gas can be converted at a high temperature into hydrogen, CO, CO2(syngas) in a steam reformer or partial-oxidation reactor, or autothermal reformer which is acombination of the first two. Most of the CO in the syngas is typically converted into carbondioxide at a lower temperature in a water-gas shift reactor. The remaining small amount of COmust be removed to below 10 ppm level. This can be done using adsorption, or membraneseparation, or catalytic preferential oxidation (at about 90°C with an air stream), or other practicalmeans. Also, there are designs with membrane reformers in the literature.

Natural gas composition and pressure: use that available at the sight of your plant. If no data can befound, use the data below:

organic sulfur 2 ppm

5 barg

References

Chemical Engineering, July 2001, pp. 37-41

AIChE Journal, July 2001, perspectives article.

A-II.10.2 Combined Cycle Power Generation

(William B. Retallick, Consultant, January 2001)

The wave of the future in power generation is the combined cycle, in which gas turbines arecombined with steam turbines, with the hot exhaust from the gas turbine used to generate steam.The combined cycle is a cascade of heat engines operating over temperatures from 1200-1300°C toabout 30°C. This broad temperature range renders the combined cycle efficient.

A gas turbine is comprised of three main parts. The compressor compresses the inlet air to thepressure in the combustor, with fuel injected into the latter. Hot combustion gases are expanded in

CD-A-II-72the turbine, which drives the compressor, with the bulk of the power produced by the turbineconsumed by the compressor.

The final stage(s) in the turbine comprise a “free” turbine. These stages are mounted on the shaftof an electric generator, rather than the shaft of the main turbine. To generate 60-cycle power, thefree turbine rotates at 3,600 RPM. Note that the main turbine rotates at 8,500 RPM. The exhaustgas from the free turbine is sent to the steam generator.

There are two kinds of steam turbines. In a condensing turbine, the exhaust steam is condensedusing cooling water, at a pressure determined by the temperature of the cooling water. In anextraction turbine, the exhaust steam is not condensed, rather it is exhausted at an intermediatepressure to be reheated and used in a condensing turbine.

The efficiency of the combined cycle is determined almost entirely by four parameters:

• The temperature of the combustion gas entering the turbine, here

assumed to be 1,250°C. • The efficiency of the compressor, here assumed to be 89%. • The efficiency of the turbine, here assumed to be 90%. • The efficiency of a steam turbine, here assumed to be 89%.

The gas turbine is operated at a compression ratio that maximizes the work produced per weight ofair, as explained in the first reference. Your turbine is equivalent or similar to the W501G turbinedescribed in the second reference. Its electrical power output is 230 MW, which fixes the size ofyour combined cycle. You are to configure a set of steam generators and steam turbines thatprovides the economic-optimum amount of electric power from the heat in the exhaust from thefree turbine.

Design Basis

The fuel gas composition (mol %) is:

Methane 96 Ethane 3 Propane 1

The gas is delivered at 400 psig.

The ambient air is 25°C and 40% relative humidity. Cooling water for the condensing turbine is at 30°C. The efficiency of an electric generator is 98%. The plant is located in Pennsylvania.

CD-A-II-73Your report should include:

The plant efficiency, kwh of fuel per kwh of electrical power.

A graph of the investor’s rate of return (IRR) as it varies with the selling price of power, for different costs for the fuel gas.

References

Chem. Eng. Prog., May, 2000, page 69.

Diesel and Gas Turbine Worldwide, July-August, 2000, page 42.

A-II.10.3 Production of Low-Sulfur Diesel Fuel

(Matthew J. Quale, Mobil Technology Company, January 2000)

There is a trend in recent environmental legislation to lower sulfur specifications in both gasolineand diesel fuels. You work for a refinery in the Delaware Valley that anticipates a new dieselspecification requiring an order of magnitude lower sulfur than currently allowed. In fact,legislation is already in the works in Europe to lower the allowable sulfur to this new level by theyear 2005. To achieve these low sulfur levels, you are to design a new catalytic hydro-desulfurization (HDS) system. This type of reactor has been in use in industry for a long time, butnever for such severe service.

This unit will require just two feeds: a liquid feed blend from your refinery, and hydrogen. Sinceyour refinery does not have a reforming unit (common hydrogen source within a refinery) or ahydrogen plant, you will have to buy the necessary hydrogen from a third party. Fortunately, agroup similar to yours (1998/1999 Penn Senior Design Group – Khandker et al., 1999) recentlydesigned a new hydrogen plant for the Delaware Valley that should be on-stream shortly beforeyour unit and they are looking for new long-term supply agreements. Use the conclusions fromtheir published report for information on the hydrogen purity and price. (I would recommendresearching current contract hydrogen prices to ensure they are charging a reasonable price,however.)

Your R&D department has done substantial pilot plant work on this new process and hasdetermined the following correlations to assist you in designing the HDS reactor. You also haveprocessing data available from an older HDS unit within your company to use as a baseline.

Processing Conditions

A common value to track while designing a HDS unit is the percent hydrodesulfurization (%HDS):

 ( wt% S) product  % HDS = 1 −  × 100%  ( wt% S) feed 

Different catalysts have different intrinsic activities, aging rates, and processingabilities/robustness. For your particular feedstock, the R&D department found the following

CD-A-II-74correlations based on a reference catalyst. Terms denoted with a “0” are the baseline data providedin Table 2. The start-of-cycle (SOC) temperature is given by

where TSOC and T0 are in °F. The aging rate is given by

The values for constants (A – F) and the base and proposed operating parameters are given in thefollowing tables. Please note that the equation and constants are Mobil Corporation internalnumbers and should be cited as such.

CD-A-II-75 The circulation and pressure values for the proposed operation are given as minimums to achieve the necessary product specifications. Increasing these values will improve the catalyst life, but result in higher capital and operating costs. It is left to you to determine the optimum values from an economic standpoint. To determine the catalyst cycle length, take 750°F as the maximum average bed temperature because higher values will produce product which has a color greater than ASTM 2.0 (the current spec).

Feed and catalyst information along with a brief overview of catalytic hydroprocessing (HDS in particular) will be presented to your group prior to beginning the project.

U.S. Patent 5,011,593 to Mobil Oil Corporation

U.S. Patent 5,474,670 to Exxon Research and Engineering

U.S. Patent 5,454,933 to Exxon Research and Engineering

A-II.10.4 Waste Fuel Upgrading to Acetone and Isopropanol

(Robert Nedwick and Leonard A. Fabiano, ARCO Chemical, January 1997)

Your plant produces two byproduct streams from an existing process. Currently, these streams arebeing sent to the on-site steam boiler where they are burned to produce high-pressure steam for thecomplex. A recent change in the environmental regulations have put the major components ofthese streams (acetone and isopropanol) on an environmental listing, which will require you tospend capital to upgrade the existing boiler and storage tanks if you continue burning. You havebeen asked to determine the optimum disposition of these streams among the following options:

1. Continue burning these streams. The capital required to upgrade the boiler and storage tanks is $10.0 MM.

2. Build a unit to produce specification grade acetone product.

3. Pay to have the streams taken off site for proper disposal at the rate of 5.0¢/lb.

Heating Value, BTU/lb (water free) 12,000 12,000

Acetone Product Specification

Acetone 99.5 wt% min

The isopropanol can be dehydrogenated to acetone per the referenced patents. The reaction ishighly endothermic and at 90% conversion does result in coking of the catalyst, which requires aone-week regeneration burnout every two months. At 80% conversion, the catalyst run length canbe extended to six months. Expected catalyst life is four regenerations after which the catalystmust be replaced at $6/lb.

Purification of the acetone product will require overcoming some azeotropes, acetone/methanolbeing the most important.

Your plant would be situated in an existing complex where much of the infrastructure is available.

The following is the situation with the Outside Battery Limit (OBL) components:

Site Development: Everything provided except:

• Site Clearing/Prep ($500M). • Control Room Upgrade ($250M).

CD-A-II-77Utilities • Firewater/N2/Air/demineralized water/potable water are all available. • Boilers are available to provide as much as 200 M lb/hr of 600-psig steam at $4.50/Mlb. • Cooling Water is at its limit and a whole new system will be required. • Electrical power is available at 4¢/kW. • All utility and electrical tie-ins are required ($1,000 M).

Storage • One fuel tank sized for the two waste fuels for 7days is available. (The acetone and IPOH streams are currently fed to a single tank before being sent to the boiler.) • Product Storage for 14 days is required. • Two Product Day Tanks for testing product quality before sending to the larger product tank are required. • Any other new fuel, product, solvent, chemical, etc., storage associated with this process is required. • Pipe runs from IBL to Storage and Loading areas are required ($500 M). • Truck Loading upgrade is required ($350 M).

Environmental • Waste water to the bio-pond can be treated for $2.00/Mgal. • The flare system can handle 150,000 lb/hr extra load upon CW failure. If higher, an additional flare will be required. • Tie-ins to these systems are required ($150 M).

Of the three options identified, only the acetone recovery has the potential for positive returns, butat the highest capital. The projected price of acetone is 25¢/lb in 2000, the target startup date forthis unit. However, acetone has experienced large pricing swings being as low as 15¢/lb and ashigh as 30¢/lb for extended periods. Acetone is also a new product line for the company and thereis some reluctance on management’s part to get involved. Your company’s philosophy is that aproject must achieve a minimum economic hurdle rate of 12% investor’s rate of return (IRR).

“Dehydrogenation Catalyst,” U.S. Patent 2,549,844 (April 24, 1951).

It is now January 1993, and the public is perceiving that the United States is burying itself in solidwaste materials. ConAgra, Inc., has approached DuPont with a proposal for a joint venture toproduce lactic acid for conversion to biodegradable polylactide plastics to be used in packaging andother markets that might help to alleviate the solid waste problem.

Under the proposal, the United States Ecological Chemical Products Company (Ecochem) willbuild a 20-million-pound acid plant based on cheese whey as a raw material at Adell, WI, wherethe Adell Whey Company will collect whey from producers within a 100-mile radius and supply itto the lactic acid plant via pipeline.

Key to the proposal appears to be the use of new technology being developed at the Oak RidgeNational Laboratory under Dr. Brian Davison. The new process is based on a three-phase,biparticle, fluidized-bed bioreactor, in which lactic acid, produced continuously in a fluidized bedof immobilized Lactobacillus delbreuckii, is simultaneously adsorbed onto a solidpolyvinylpyridine resin moving countercurrent to the fermenter beer. In this way, the pH can bemaintained at the optimum 5.5 and product inhibition of the fermentation is minimized. As aresult, fermentation rates have been increased 4- to 10-fold higher than the conventionalfermentation process and the acid product can be recovered by methanol extraction.

As the Planning Manager for Chemicals, you have been asked to evaluate the techno-economics ofthe proposal as compared with the alternative conventional fermentation process and advise theExecutive Committee of the financial expectations for the venture. It appears that if the design andfinancial evaluation can be completed by May 1993, the plant can be constructed for start-up inJanuary 1996.

Potera, C., Genetic Eng. News, 1 (November 1, 1992).

A-II.10.6 Ethanol for Gasoline from Corn Syrup

In its environmental program, the Bush administration is evaluating clean-burning, low-volatile

fuels for automobiles. One alternative is to convert a farm product, corn syrup, to motor-gradeethanol.

Your consulting company is requested to design a 100,000-metric tons/year, automobile-grade

ethanol plant using corn syrup as the feedstock. After designing the process and determining itstotal cost, the price subsidy to make ethanol competitive with current gasoline prices should bedetermined.

In creating your design, give special consideration to processes that reduce the energy expenditureof the plant. In one such process, pervaporation membranes are used to dehydrate ethanol.Pervaporation is a membrane separation process in which the feed and residue streams are liquid,but the permeate is a vapor. The combination of permeation and evaporation in the membranegives rise to separation factors much greater than can be accomplished by distillation and can beused to break azeotropes.

Prepared by James R. Phimister

University of Pennsylvania Department of Chemical Engineering Philadelphia, PA 19104/6393

The General Algebraic Modeling System (GAMS) is a programming language that provides aflexible framework for formulating and solving linear, nonlinear and mixed-integer optimizationproblems. Among other attributes, its syntax allows for declaring associations among variables,constants, and constraints in the form of sets. Through this syntax, input files are writtencompactly and similarly to the typical formulations of optimization problems. In addition, GAMSprovides a wide array of solvers to optimize a variety of problem formulations including linearprograms (LPs), nonlinear programs (NLPs), mixed-integer linear programs (MILPs, but referredto as MIPs by GAMS), and mixed-integer nonlinear programs (MINLPs). Here, it is possible only to provide a limited overview of the capabilities of GAMS.Examples discussed in the text are presented to illustrate the structure of input files for an LP, anNLP, and an MIP. These input files can be copied from the CD-ROM that accompanies thisbook. It is recommended that the reader copy and run the GAMS input files, and observe theresults. The files can be modified and rerun to observe how the optimal solutions change. Aswith many software applications, one of the best means for learning GAMS is through hands-onexperience. For interested readers a detailed presentation of GAMS is provided in GAMS: A User’sGuide: Release 2.25 (Brook, A., D. Kendrick, and A. Meeraus, Scientific Press, San Francisco,1992). The solvers available in GAMS are presented in GAMS - The Solver Manuals (GAMSDevelopment Corporation, Washington, DC, 1996). In its simplest form, GAMS operates on a user-supplied input file (normally denoted witha .gms or .number extension to the filename), which encodes the mathematical formulation ofthe optimization problem being examined. Selection of the word processor for use in editing theinput file is left to the discretion of the user. Note, however, that files should be saved in ASCII

CD-G-1or TEXT format. Thus, the Notepad word processor included in the Windows operating systemis a good choice. GAMS is not likely to be able to run files saved in another format (e.g. that of.doc files in WORD) since the default save options of many word processors add formattingcodes to the saved file. GAMS operates on a variety of platforms, with execution of the program initiated fromthe command prompt line. GAMS is run using the executable (GAMS) followed by the name ofthe input file at the command prompt, that is:

GAMS filename.extension

In the WINDOWS operating system, GAMS is run through the Run window, shown in Figure 1,which is obtained by left clicking the Start window and then clicking Run. In this case, theexecutable (GAMS) and the input file (CASC.1) are located in the directory C:\GAMS. Theoutput file, CASC.lst, is placed in the same directory as the GAMS executable, not the directoryof the input file.

Figure 1. GAMS run from WINDOWS or NT.

GAMS operates on an input file in two stages:

1. Compilation. This stage ensures that the input file is understood by GAMS. The compiler checks for errors in the input file, ensuring that the file abides to a specific format, does not contain syntax errors, and uses an appropriate solver. The compiler does not solve the problem or indicate that a solution exists. When the compiler locates errors in the input file, the errors are flagged and written in the output file (e.g.,

CD-G-2 CASC.lst) before GAMS terminates. The user must then correct the input file. When the compilation is successful, GAMS proceeds to stage 2.

2. Execution. With the input file readable, GAMS proceeds to carry out the optimization using an appropriate solver for the problem formulation (e.g., LP, NLP, and MIP). Note that the solver declared by the user must be applicable to the formulation. For instance, an LP solver cannot be used to solve an NLP. GAMS writes to the output file, providing information on whether the solution was obtained, and if so, the solution values. Output can be controlled using display options in the input file.

where Qsteam, Qcw, and R1, R2, R3, R4 are all non-negative real numbers. The GAMS input file,CASC.1, in Figure 2 contains this LP formulation. Note that it closely resembles the writtenproblem formulation and is equivalent to the input file in Example 10.4.

MODEL CASCADE /ALL/;

The output file, CASC.lst, contains a wealth of information. Of particular interest is the SolveSummary:

S O L V E S U M M A R Y

MODEL CASCADE OBJECTIVE Z

TYPE LP DIRECTION MINIMIZE SOLVER MINOS5 FROM LINE 38

**** SOLVER STATUS 1 NORMAL COMPLETION

**** MODEL STATUS 1 OPTIMAL **** OBJECTIVE VALUE 50.0000

that shows normal completion of the linear program (LP) by the solver MINOS5 to one optimalsolution; that is, Z = 50.0.

Statements

Although much of the input file, CASC.1, is self-explanatory, it is important to understand thestructure of a GAMS input file and its statements. In its simplest form, an input file must consistof statements for:

CD-G-4 1. Variable declarations and assignments 2. Equation declarations 3. Equation definitions 4. A model declaration, with an appropriate solve statementInitially, it is recommended that all statements be ended with a semicolon, as statements withoutsemicolons may cause compiler errors.

Variable Declaration Statement. Each variable must be declared in the input file. Note that avariable for the objective function must also be declared. In Figure 2, the objective functionvariable is denoted as Z.

Equation Declaration Statement. Each constraint as well as the objective function must bedefined with a name.

Equation Definition Statement. For every declared equation name, a corresponding equationmust be defined. To define the equation: 1. Every equation name is restated as written in the declaration followed by two periods and at least one space. 2. The equation is stated using the declared variables and constants, the operators (+, -, ≥, etc.) and GAMS functions (sine, sum, etc.). 3. Each equation is defined as a statement, and hence, is ended with a semi-colon.

The following relational operators are defined:

Relation Syntax Equality constraint (=) =E= Less than or equal to (≤) =L= Greater than or equal to (≥) =G=

Note there is no definition for the strict inequalities less than (<) or greater to (>). This omissionis intentional and does not result in any loss of generality by GAMS. The most commonly used arithmetic operators are:

In addition, several built-in functions are available, including:

Model Declaration Statement. The model declaration statement defines a user-specified namefor the model and declares the equations to be included in the model. For novice GAMS users, itis recommend that all equations be included. Hence, the model declaration statement is:

MODEL MODEL_NAME /ALL/;

Solve Statement. The solve statement defines the:

1. Model to be optimized (defined previously in the model declaration statement). 2. Type of solving procedure (LP, MIP, etc.). 3. Type of optimization; that is, minimization or maximization. 4. Variable to be optimized.It has the form:

SOLVE MODEL_NAME USING PROBLEM_TYPE MINIMIZING OBJEC_FUNC_VARIABLE;

The key words ‘SOLVE’ and ‘USING’ must be present in the statement, as well as either‘MINIMIZING’ or ‘MAXIMIZING’

CD-G-6 In Figure 2, the model name is CASCADE, the type of solving procedure is LP since onlylinear constraints and a linear objective function appear. The objective function variable, Z, isminimized, and consequently, the solve statement is:

SOLVE CASCADE USING LP MINIMIZING Z;

There are several optimization formulations that can be solved, the most common being:

Formulation GAMS Syntax

2. EXPANDED FEATURES - DOCUMENTATION, VARIABLE

REDECLARATION, AND DISPLAY

Using the basic features of GAMS, a large array of optimization problems can be solved. Thereare, however, several features that greatly improve the ease of formulation and the readability ofthe input and output files.

Syntax. The GAMS compiler does not distinguish between upper and lower case characters.Hence, to allow for more readable text, both cases may be used interchangeably in an input file.

Documentation. It is important to add documentation to the input file to simplify debugging as

well as to clarify the formulation of the optimization problem. Documentation can be interspersedthroughout the file by placing an asterisk (*) in the first column of a documentation line. Notethat the asterisk alerts the compiler to overlook the line. Documentation can also be includedwithin declaration statements. Any character string placed after a variable or equation has beendeclared, but prior to a comma or semi-colon, is considered to be documentation attributed to thevariable or equation. These character strings also appear in the output file.

CD-G-7Variable Redeclaration. It is often desirable to re-declare a variable with bounds provided. Forexample, variables w and x are re-declared such that bounds greater than or equal to zero areapplied using the statement

POSITIVE VARIABLES w,x;

Alternatively, for real-valued variables, this can be accomplished by defining equations thatincorporate the bounds on the variables. However, the use of re-declared variables provides for amore concise input file. The key words to define the bounds on a variable are

GAMS Syntax Range on Variable

Note that binary and integer variables must be re-declared for mixed-integer programming. Todefine the binary variable, y, this is accomplished using the statement:

BINARY VARIABLE y;

Variable Display. After computing a solution, GAMS displays the variable values in the outputfile. In addition, GAMS can redisplay values in tabular form toward the end of the output file.To accomplish this, a DISPLAY statement is added following the MODEL statement. While arange of possible outputs can be displayed, the level (or final) outputs (denoted with a ‘.L’extension to the variable) are often of most concern. For example, the final values for R1, …, R4,Qcw, and Qs, are requested in tabulated form in the output file, using the input statement:

DISPLAY R1.L,R2.L,R3.L,R4.L,Qcw.L,Qs.L

Subsequently, these are displayed in the output, as shown in Figure 3.

Figure 3. Displayed values requested in input file CASC.1

Bounds and Initial Conditions. When the bounds using variable re-declaration statements areinadequate, additional statements are available. To set a lower bound for x, use:

x.lo = 10.0;

and to set an upper bound for x, use:

x.up = 100.0;

An initial starting point for the solver is supplied by the user with the level (‘.l’) extension:

x.1 = 50.0;

Bounding and initialization statements appear after the variable declaration statements, but priorto the equation declaration statements. For nonlinear programs (NLPs) it is important to provide bounds and initial conditionswhere possible. This is often necessary because:1. It is often difficult for a nonlinear solver to locate a feasible solution (one that satisfies the constraints), especially when the initial guessed values are poor. When the user provides a feasible starting point, the likelihood of successful convergence to an optimal solution is greatly improved.2. Nonlinear programs are often multi-modal; that is, the surface of the objective function contains numerous local minima. With physical insights, the user may be able discern where the global minimum is likely to exist, and hence, provide initial conditions near this point. Likewise, the user may be able to discern where a solution cannot exist, and hence provide tighter bounds on the search space. GAMS does not seek all local minima.3. GAMS defaults initial variable values to zero. When a variable appears alone in a denominator, or is the argument of a log function, the solver may abort.

CD-G-9 * This is the input file for Example 10.4 * The GAMS file formulates an LP to be solved for minimum utilities

* Declare the variables

* Note that Z (used for the objective function) * 1. must be declared * 2. should not be made positive

GAMS has a powerful feature which allows sets to be declared. Sets allow for subscriptedvariables used in variable and constant declarations, as well as equation definitions. As anexample of the utility of declaring sets, an optimization problem might contain the fiveconstraints:

xi ≤ 10 ∀i, where i∈{1,2,…, 5}

Likewise, by employing sets in GAMS, input files can be written with subscripted constraintswritten in a single line. Note any input file that incorporates sets can be written in GAMSwithout defining sets. However, input files with defined sets can be written concisely, and areoften easier to debug and update. To illustrate the definition of sets and how they are employed, Example 10.8 isconsidered. In this example, it is desired to determine the least number of matches between hotand cold streams while providing maximum heat recovery. Note that the problem has beenrestated in Figure 6. The GAMS input file, MATCH.1, is shown in Figure 7, and provided on theaccompanying CD-ROM.

Set Declarations. In Example 10.8, variables are defined with reference to whether there is amatch between a hot stream and a cold stream, and if a match exists, over which interval in theheat cascade. Hence, yH2,C2 is a binary variable representing whether hot stream 2 is matched tocold stream 2, and QH1,C1,3 is a continuous variable that represents the heat transferred betweenhot stream 1 and cold stream 1 in interval 3. In formulating the input file, H2 is a member of the set of hot streams, termed HOT,where: HOT= {H1, H2, S}

CSIDE(COLD,INT).. SUM(HOT,Q(HOT,COLD,INT)) =E= QC(COLD,INT);

MODEL MATCH /ALL/;

Figure 7. MIP input file for matching heat exchangers.

CD-G-14Likewise, a set COLD is defined as the cold streams {C1, C2, W} and the set INT = {1,2,3,4,5}defines the intervals in the heat cascade. These sets are defined by the statement: SETS

HOT hot streams / H1, H2, S /

COLD cold streams / C1, C2, W / INT interval / 1,2,3,4,5 /;

Note that the elements in the set are defined within the slashes (/). The set name is the first wordin the line. Any character strings following the set name and preceding the first slash aredocumentation.

Data Statements: Scalars, Parameters and Tables. It is often desired to declare constants,which can be referred to in the equation definitions. For instance, a constant UPPER is declaredand assigned the value 100.0 using the GAMS statement:

Scalar UPPER defines an upper bound /100.0/;

Similarly an array of scalars, termed a parameter can be declared. Parameters are defined for setswhose elements are the names in the array. For example, when the feed temperatures for the setof cold streams {C1, C2, W} are 60, 116, and 38°C, respectively, a parameter TENTER isdeclared using the statement:

Parameter TENTER entering cold stream temp /60, 116, 38/;

Tables have a dimension of two or greater and declare inter-related parameters. In

Example 10.8, a variable U is declared. On inspection, U can be declared as a 2-dimensionaltable whose elements are entered in GAMS using the TABLE statement:

TABLE U(HOT,COLD) upper bound on heat transfer between streams

CD-G-15The format for the table declaration is flexible; spacing between the elements is greater or equalto one space. It is recommended that the TAB key not be used when constructing tables as texteditors save tab spaces differently, often creating parsing problems for the compiler. When tablecells are left blank, they are assigned the default value of zero.

Variable Declarations. Variables can be declared over sets (that is, with sets as their elements).In Example 10.8, the variable Ri,k is declared R(HOT,INT), Qi,j,k is declaredQ(HOT,COLD,INT), and the binary variable yi,j is declared Y(HOT,COLD).

Equation Declarations. Sets allow equations to be declared for all (∀) elements in a set. Forexample, rather than declare nine equations and define each equation to express the maximumheat transferred between its match (H1-C1, H1-C2, H1-W, H2-C1, etc.), the set of equations,QMATCH(HOT,COLD) can be declared, and subsequently, defined to encompass all matchesbetween the elements in HOT and COLD.

Equation Definitions. The first term in an equation definition statement is the equation name.Hence, when the declaration specifies that the equation utilizes sets, the equation definition mustdo so as well. This is shown for the set of equations, QMATCH(HOT,COLD), with an equationdefined to provide an upper bound on heat transferred in a match:

When it is necessary to refer to a specific element in a set, quotation marks are used. Forexample, the following statement defines constraint S in Figure 7:

S.. R("S","1")+Q("S","C2","1") =E= QH("S","1");

which defines the heat balance for the stream of steam in interval 1. A number of set specific functions are available in GAMS. The most commonly usedfunction is SUM, which sums over all of the elements in a given set. This is shown in theobjective function Z:

OBJ.. Z =E= SUM(HOT,SUM(COLD,Y(HOT,COLD)));

The above statement sums y over all of the elements in both sets HOT and COLD.

4. DEBUGGING

Two types of errors are encountered: errors during compilation and errors during execution. Allerrors are written in the output file.

Compilation errors. Many errors are usually reported in the output file the first time an input fileis run. Since errors further down in the input file are often a result of an error earlier in the file, itis recommended that one proceed from the top of the file downwards, rerunning the input fileafter several errors have been corrected. Compilation errors indicate that the input file contains statements that are not recognizedby the compiler. Common compilation errors include: 1. Syntax errors, such as failure to end a statement with a semi-colon. 2. Left and right parentheses not matching. 3. Incorrect references to a variable name.Until the compilation errors are corrected, GAMS is unable to execute the model.

CD-G-17 During compilation, GAMS copies the input file to the output file, adding row numbers inthe left-hand column. When compilation errors are encountered, a dollar sign ($) followed by anerror number is indicated. At the end of the file, a brief description of the errors is provided. Consider the input file CASC.1 in Figure 2 with the following modifications: 1. Qs is omitted in the variable declaration statement. 2. The semi-colon following the statement defining equation B1 is omitted.In the corresponding output file, shown in Figure 8, three errors are flagged. The flags arenormally directly below the statement at which GAMS anticipates that the error occurs. In thiscase, GAMS detects and flags the undeclared variable Qs. However, the error occurs because QSis omitted from the variable declaration. In addition, because the semi-colon is omitted, GAMSis unable to parse statement B1 from statement B2. Finally, GAMS does not check the solvestatement because the other errors are identified.

Execution errors. Execution errors occur when the program compiles successfully and a solveris attempting to locate the optimum solution. It is usually more difficult to correct these errorsbecause they are related to the optimization algorithm. When trying to resolve execution errors, itis recommended that the user:1. Check the optimization formulation and its transcription into the GAMS input file. Is the formulation correct? Has the formulation been copied correctly to the input file? Are all of the variables and equations that use sets correctly stated?2. Consider adding or altering the bounds.3. When the solver performed an illegal operation, such as a divide by zero, check whether the equations can be modified/transformed to avoid this error. Inspect whether different initial conditions or stricter bounds avoid premature termination of the algorithm.4. When the solver reports the problem is infeasible, check whether the problem is an LP or an MILP. If so, the formulation is likely to be inconsistent and the model should be re- examined. If not, consider slackening or removing some of the constraints, and re-running GAMS until a feasible solution is obtained. Then, re-introduce constraints that have been removed or slackened, attempting to discern why infeasibility is reported. Also, for an NLP, the solver may not be able to locate a feasible solution. Examine equations shown to be

CD-G-18 infeasible in the output file. Check whether these equations can be modified or removed to improve the likelihood of convergence. Try to provide an initial feasible solution.

**** 3 ERROR(S) 0 WARNING(S)

It may be useful to observe the level (final) values for specific variables after the firstiteration. To stop the solver after one iteration, the following statement is added to the input file,immediately before the solve statement:

OPTION iterlim = 1

In addition, after the solve statement, add appropriate display statements for those variables to betabulated in the output file.

3. Equipment items are sized and re-sized when modified.

4. Capital costs, operating costs, and the total investment are evaluated for a project.

5. Results are presented to be reviewed, with modifications as necessary and re-

evaluation.

CD-16.7-1 Aspen IPE begins with the results of a simulation using one of the major processsimulators. The program accepts results from ASPEN PLUS, HYSYS.Plant,CHEMCAD, PRO/II, and other simulators. To estimate equipment sizes and costs, it isnecessary to prepare simulation results for loading into Aspen IPE. This is normallyaccomplished by augmenting the simulation file in two ways. First, to estimateequipment sizes, Aspen IPE usually requires estimates of mixture properties not neededfor the material and energy balances, and phase equilibria calculations performed by theprocess simulators. For this reason, it is necessary to augment the simulation report fileswith estimates by the simulator of mixture properties, such as viscosity, thermalconductivity, and surface tension for each of the streams in the simulation flowsheet.Second, Aspen IPE estimates equipment sizes using the simulation results computed bythe more rigorous, rather than approximate, simulation subroutines. Consequently, whenthe approximate DISTL and RSTOIC subroutines are used in ASPEN PLUS, these mustbe replaced by more rigorous subroutines, such as the RADFRAC and RPLUGsubroutines. This replacement can be viewed as the first step in computing equipmentsizes and costs.

After the simulation file is augmented, the revised simulation is run and theresults are sent to Aspen IPE. Note that the ASPEN PLUS and HYSYS.Plant simulatorscontain menu entries to direct the results to Aspen IPE. For details, the reader is referredto course notes prepared at the University of Pennsylvania (Nathanson and Seider, 2003),which are provided in the file, Aspen IPE Course Notes.pdf, on this CD-ROM. Thissection presents estimates of equipment sizes and purchase and installation costs usingAspen IPE for two examples involving: (1) the depropanizer distillation tower presentedon the CD-ROM (either HYSYS → Separations → Distillation or ASPEN PLUS →Separations → Distillation), and (2) the monochlorobenzene (MCB) separation processintroduced in Section 4.4, with simulation results using ASPEN PLUS provided on theCD-ROM (ASPEN → Principles of Flowsheet Simulation → Interpretation of Input andOutput → Sample Problem). Just the key specifications and results are presented here.The details of using Aspen IPE for these two examples are presented in the file, AspenIPE Course Notes.pdf.

CD-16.7-2Example 16.18 Depropanizer

The depropanizer distillation tower in Figure 16.16 is designed and simulated

using the procedures described on the CD-ROM (either HYSYS → Separations →Distillation or ASPEN → Separations → Distillation). In summary, for the pressuresshown, using the DSTWU subroutine for the specification R = 1.75Rmin, the reflux ratio,number of equilibrium stages, and the feed stage are estimated to be: R = 6.06, N = 14,and NFeed = 7. When the tower is simulated with these specifications and D/F = 0.226, toachieve the desired distillate purity, the RADFRAC subroutine adjusts the reflux ratio to8.88. In this example, it is desired to estimate the total permanent investment, CTPI, usingAspen IPE. The material of construction throughout is carbon steel.

Figure 16.16 Specifications for design of the depropanizer

distillation tower

CD-16.7-3SOLUTION

For the depropanizer system, Aspen IPE performs mechanical designs, andestimates sizes, purchase costs, and associated installation materials and labor costs forthe distillation tower, condenser, reflux accumulator, reflux pump, and reboiler. Thedesigner can add a reboiler pump (to pump liquid from the sump to the reboiler), as wasdone in obtaining this solution. Aspen IPE uses many parameters to estimate equipmentsizes and to specify the characteristics of utilities, with default values built in that can bereplaced by user-specified values. Particular attention should be paid to the IPE DesignBasis parameters, such as the design pressure and temperature, the overdesignallowances, the residence times in the process vessels, and the tower specifications. Forthe depropanizer complex, a few changes were made to the default parameters, includingthe tray efficiency (0.8), bottom sump height (10 ft), and vapor disengagement height(above the top tray, 4 ft). The other default parameters are listed in Appendix II of thefile, Aspen IPE Course Notes.pdf, on the CD-ROM.

For the condenser, Aspen IPE uses the cooling water utility. However, its defaultinlet and outlet temperatures were changed from 75 and 95°F to 90 and 120°F. Also,Aspen IPE has three built-in utilities for steam at 100, 165, and 400 psia. Because 100psia steam condenses at 377.8°F and the bubble point temperature of the bottoms productat 252 psia is 260.8°F, when 100 psia steam is used in the reboiler, ∆T = 117°F, whichoften results in undesirable film boiling as discussed in Section 13.1 of the book. Toreduce the approach temperature difference, and assure nucleate boiling, a low pressuresteam utility, at 50 psia, is defined.

After the parameters for estimating equipment sizes and the utility parameters areadjusted, and a new steam utility is defined, the simulation units (blocks, modules, orsubroutines) are mapped into Aspen IPE. In this case, there is only one distillation unit,D1, to be mapped. The default mapping results in: (1) a tray tower, (2) a shell-and-tubeheat exchanger with a fixed tube sheet for the condenser, (3) a horizontal drum for thereflux accumulator, (4) a centrifugal reflux pump, and (5) a kettle reboiler with U tubes.

CD-16.7-4To use a kettle reboiler with a floating head, or one of the other built-in reboilers, thedefault mapping is deleted and replaced with the preferred mapping. Similarly, thedefault mapping for the condenser, a shell-and-tube heat exchanger with a fixed tubesheet, can be replaced with a shell-and-tube heat exchanger having a floating head.When the mapping for the simulation unit, D1, is completed, sizes have been estimatedby Aspen IPE for all of the equipment items. Note that for this distillation complex areboiler pump is added by the designer and mapped separately by Aspen IPE. Note alsothat the equipment sizes can be adjusted by the designer before Aspen IPE estimatesequipment costs, although no adjustments have been made here for the distillationsystem.

In the next step, Aspen IPE estimates the purchase and installation costs. Beforeproceeding, the designer can (1) apply one of six engineering contractor profiles, whichdetermine the engineering execution procedure, and (2) adjust the standard basis, whichdefines the nature of the site and workforce. Here, the default values may correspond toinappropriate costs for the following reasons. When designing small plants, the PlantEngineer or Local Contractor profiles are preferable. For this distillation system, whichis the only system in the plant, it is important to replace the default project type (grassroots/clear field) with plant addition – suppressed infrastructure. The latter instructsAspen IPE to omit items involving electrical switchgear and transformers, which are notneeded when adding this distillation system to an existing process facility. After thestandard basis has been adjusted, Aspen IPE evaluates all of the equipment items in theproject. During the evaluation, purchase and installation costs are estimated. For thispurpose, Aspen IPE utilizes design, work-item, and cost models that have been developedand updated annually, in accordance with industry design codes and costs for numerousprocess plants, since the mid-1970s. Given the broad spectrum of Aspen IPE usersworldwide, Aspen IPE purchase cost estimates are based upon an extensive data base ofmaterial and construction labor costs and detailed, though preliminary, design methods.

For installation costs, Aspen IPE does not use bare module factors as discussed inSection 16.3 of the book. Rather, rigorous methods are used to estimate the costs of

CD-16.7-5materials, labor, and construction equipment. These methods are based upon detaileddesign calculations for foundations, platforms, piping, instrumentation, electricalconnections, insulation, and painting, among other items involved in the installation. Forexample, for concrete foundations, the dimensions of the foundation and the amount ofconcrete are estimated based upon the height and weight of the tower, soil conditions,wind velocity, and seismic zone. For piping and instrumentation types, quantities, andsizes, Aspen IPE uses self-contained, user-adjustable, P&ID templates that are unique toeach type of equipment. Aspen IPE uses its library of piping and instrumentation models,mechanical design methods, and equipment and stream information, to develop lists ofmaterials for piping and instruments, with associated material costs and installationhours. Consequently, the installation cost estimates by Aspen IPE are more accurate thanthose obtained using bare-module or factored-cost methods.

For the six equipment items in the depropanizer distillation system, including theadded reboiler pump, the key equipment sizes and cost estimates are shown in Table16.33. Note that Aspen IPE designed the condenser to be a shell-and-tube heat exchangerwith two parallel units, each having two tube passes and a correction factor, FT = 0.64. Itshould be possible to improve this design by re-sizing the unit to obtain a correctionfactor close to unity, eliminating one of the parallel units. Figure 16.17 shows moredetails for the tray tower from the Capital Estimate Report. For details of the otherequipment items, see Appendix III of the file, Aspen IPE Course Notes.pdf, on the CD-ROM. Also, these results can be reproduced by accessing the DEC3RP folder (on theCD-ROM in the Aspen Eng. Suite folder) from within Aspen IPE. Note that the DEC3folder does not include the reboiler pump. The calculations were carried out using AspenIPE, Version 11.1, with the design and cost basis date being the First Quarter 2000.

CD-16.7-6Total Permanent Investment

Aspen IPE also computes the total permanent investment, CTPI, as defined inTable 16.9 of the book. However, here, the total permanent investment is computed bythe spreadsheet, Profitability Analysis-1.0.xls, which is discussed in Section 17.8. Whenusing the Aspen IPE option in the spreadsheet, the user enters the following values,which are obtained from Aspen IPE:

Total Direct Materials and Labor Costs $757,500

Note that the total direct materials and labor costs, $757,500, includes items notchargeable to the individual equipment items in Table 16.33. For the details of obtainingthese values from Aspen IPE, see the file, Aspen IPE Course Notes.pdf, on the CD-ROM.

Figure 16.17 Estimates of equipment sizes and purchase

and installation costs for the depropanizer tray tower.

CD-16.7-9Example 16.19 Monochlorobenzene (MCB) Separation Process

The monochlorobenzene (MCB) separation process in Figure 16.18 is designed

and simulated using the procedures described in Section 4.4 and on the CD-ROM(ASPEN PLUS → Principles of Flowsheet Simulation → Interpretation of Input andOutput → Sample Problem). In this example, it is desired to estimate the total permanentinvestment, CTPI, using Aspen IPE.

Figure 16.18 Process flowsheet for the MCB separation

process.

CD-16.7-10SOLUTION

The simulation results were computed initially using the DISTL subroutine inASPEN PLUS. When this is replaced by the RADFRAC subroutine, prior to usingAspen IPE, the reflux ratio is adjusted from 4.29 to 3.35 and the stream flow rates differslightly (< 1%).

Because the absorber has a tray efficiency of 20%, while the tray efficiency of thedistillation column is 60%, the two towers must be mapped separately. Also, the heatexchanger, H1, is too small to be mapped as floating-head, shell-and-tube heat exchanger.Consequently, this mapping is replaced by a double-pipe heat exchanger. Finally, theunits, M1, S1, and T1, are mapped as Quoted Items having zero cost by Aspen IPE.

After the mapping and sizing are completed (i.e., the equipment sizes arecomputed), as for the depropanizer in Example 16.18, the MCB separation process can beviewed as representing an addition to an existing plant. Consequently, the standard basisprofile is selected to be Local Contractor and the project type is selected as plantaddition – suppressed infrastructure. This is because a full grass roots/clear fieldinstallation would provide an unnecessary supporting power distribution substation andcontrol system equipment for this small separation plant, which typically would besupported as a neighboring facility and not built as a separate entity. After the standardbasis has been adjusted, Aspen IPE evaluates all of the equipment items in the mapping.During the evaluation, purchase and installation costs are estimated. For 11 equipmentitems, the key equipment sizes and cost estimates are shown in Table 16.34, with detailsof the equipment items provided in Appendix IV of the file, Aspen IPE Course Notes.pdf,on the CD-ROM. Also, these results can be reproduced by accessing the MCB folder (onthe CD-ROM in the Aspen Eng. Suite folder) from within Aspen IPE. The calculationswere carried out using Aspen IPE, Version 11.1, with the design and cost basis date beingthe First Quarter 2000.

CD-16.7-11Total Permanent Investment

Aspen IPE also computes the total permanent investment. However, in thistextbook, the total permanent investment is computed by the spreadsheet, ProfitabilityAnalysis-1.0.xls, which is discussed in Section 17.8. When using the Aspen IPE optionin the spreadsheet, the user enters the following values, which are obtained from AspenIPE:

Total Direct Materials and Labor Costs $785,700

Note that the total direct materials and labor costs, $785,700, includes items notchargeable to the individual equipment items in Table 16.33. For the details of obtainingthese values from Aspen IPE, see the file, Aspen IPE Course Notes.pdf, on the CD-ROM.

TOTAL $154,400 $671,900

CD-16.7-1317.8 PROFITABILITY ANALYSIS SPREADSHEET

This section shows how to use purchase and installation cost estimates fromAspen IPE, and other sources, together with an economics spreadsheet by HolgerNickisch (2003) to estimate profitability measures for the monochlorobenzene (MCB)separation process, which was introduced in Section 4.4. In Section 16.7, IPE was usedto estimate the total permanent investment for this process. The economics spreadsheet,Profitability Analysis-1.0.xls, is on the CD-ROM that accompanies this textbook.

Holger Nickisch, a graduate of the University of Pennsylvania, with dual degrees

in chemical engineering and business, designed the spreadsheet for use with Chapters 16and 17 of Product and Process Design Principles: Synthesis, Analysis, and Evaluation(Seider, Seader, and Lewin, Wiley, 2004). It replaces Version 3.0 of an earlierspreadsheet, entitled HNP.xls.

the most common sources of error when setting up a complicated spreadsheet inMicrosoft EXCEL. The use of VBA makes it possible to avoid common mistakes inentering specifications, allows the output to be formatted into presentable pages, andensures that the output is not altered inadvertently after specifications have been entered.The user of the spreadsheet is not required to know VBA.

General Instructions for Use of Profitability Analysis-1.0.xls

Depending on the version of EXCEL being used, the procedures to activate

“Macro Code” differ. In EXCEL 97, when the spreadsheet is loaded, the user is askedwhether macros should be enabled. An affirmative response is necessary, after whichEXCEL loads the complete file, which contains many worksheets most of which arehidden. In EXCEL 2000, and later versions, three security settings under the Tool,Macros menu are offered. The highest setting does not allow VBA code to be opened.

CD-17.8-1The intermediate setting causes the user to be prompted, as in EXCEL 97, and the lowestsetting causes VBA code to be opened without user approval. The latter isrecommended.

After the spreadsheet is started with the “Macro Code” activated, the introductorypage is displayed briefly. Then, the Login dialog box appears in which a user name andpassword must be entered. For students, the user name is ‘student’ and the password is‘engineer’ (which can be altered). When proper entries have been provided, theSave/New dialog box appears in which the user selects either Start New Analysis or LoadExisting Analysis and clicks on the OK button.

When a new analysis is initiated, the Step 1 dialog box is displayed, into whichinput specifications are entered. The user provides entries for the title of the process, thename of the product, the location of the plant site, and the site factor (which is obtainedfrom Table 16.13). Then the annual operating hours are entered, either hr/yr, day/yr, orthe operating factor (fraction of hours in operation per year). Finally, the timelines andinvestment distribution are provided for the total permanent investment, CTPM, and theworking capital, CWC. When the Timelines button is depressed, the Timelines dialog boxappears. Entries are provided for the starting year, the number of years of design andconstruction, and plant life in years. When the Investment Distribution is pressed, theInvestment Distribution dialog box appears, in which percentage distribution in each yearcan be adjusted for CTPM and CWC. Note that when the percentages sum to 100%, the sumis displayed with a green background. Otherwise the background is red, signaling to theuser that corrections must be made. Specific entries are shown below in Example 17.32,in which a profitability analysis for the MCB separation process is completed. Afterthese specifications are completed, the OK button is depressed and the Input Summaryform is displayed. By scrolling down on this form, the following sections appear:

Note that the specifications just entered in the General Information section are displayed.Associated with each section heading is a blue button: CLICK HERE FOR MENU.When pressed, the spreadsheet menu appears at the top of the screen. To enter thespecifications in any section, point anywhere within the section, except for the sectionheading. This produces a dialog box that guides the specification of many of the requiredinputs. Other entries are provided in dialog boxes produced by pressing the Menubuttons. Note that for most entries default values are provided. These are displayed onthe Input Summary form and remain unaltered unless new entries are provided. Thedefault values are those recommended in Chapters 16 and 17. Next, the entries in eachsection are described.

General Information

These entries have been discussed above. To produce the Step 1 dialog box, pointto the General Information section and left click.

Chronology

This section lists each year during the life of the project, beginning with thestarting year. The Action in each year is indicated as Design, Construction, or Production.As discussed above, these specifications are entered using the Step 1 dialog box, which isobtained by pointing to the Chronology section and left-clicking. To change the

CD-17.8-3Production Capacity (i.e., the percent of design capacity) in each year and the MACRStax-basis depreciation schedule, on the Menu, press the Options button, which producesthe Options dialog box. Press the Production Capacity tab to enter Production Capacityat full production (the percent of design capacity at full production) and information toramp-up to full production, that is, the years to achieve full production, and the StartProduction percentage (the percent of design capacity during the first year of production.)Note that a linear ramp is computed. Also, the number of operator shifts per week isspecified. Using the Depreciation Schedule tab, the number of years in the MACRS Tax-basis Depreciation schedule is specified. Note that 5, 7, 10, and 15 year schedules aredisplayed.

Product Information

When left-clicking within this section, the Product Units dialog box is displayed.In this box, the unit in which the primary product is specified is entered in five charactersor less (e.g., lb). This produces the Capacity and Product Price dialog box in which thecapacity of the plant is entered (e.g., lb/hr of MCB) and the product price is entered in$/unit (e.g., $/lb MCB).

Raw Materials

Similar entries are provided for each raw material. When left-clicking within theRaw Materials section, the Raw Materials dialog box is produced. To add a raw material,press the Add button. This produces the Raw Materials: NEW dialog box, in which theraw material name and unit of measure are entered (e.g., FEED and lb). This producesthe Raw Materials: ‘FEED’ dialog box, within which the units of the raw material perunit of product (e.g., lb FEED/lb MCB) and the price per unit of raw material (e.g., $/lbFEED) are entered. Subsequently, this entry can be edited by displaying the RawMaterials dialog box and selecting an existing raw material using the pull-down menu.The entries for this raw material can be edited by pressing the Edit button or the rawmaterial can be deleted by pressing the Delete button.

CD-17.8-4Equipment Costs

When left-clicking within the Equipment Costs section, the Equipment Costsdialog box is produced. Herein, equipment items are identified to be in one of fivecategories: Fabricated Equipment, CFE; Process Machinery, CPM; Spares, Cspare; Storage,Cstorage; or Catalysts, Ccatalyst; as grouped in Table 16.9 and discussed in Section 16.3.Entries in each category can be entered, edited, and deleted. After one of the categoriesis selected, a dialog box, with the name of the category, appears. In this box, the userchooses to add, edit, or delete an equipment item. When pressing the Add button, theNew Equipment Item dialog box appears. The user chooses to enter (1) the purchase costonly, (2) the purchase cost and bare module factor, or (3) the bare module cost. For thefirst option, the Entering Purchase Costs Only dialog box appears, in which theequipment name and purchase cost are entered. In this case, a default bare module factor,3.21, is used. Note that to alter the default bare module factor, on the Menu, press theOptions button to produce the Options dialog box. Select the Derived Bare ModuleFactor tab. On this form, the factors in Table 16.10 are entered, as fractions of thepurchase cost to compute the cost of installation materials, CM; labor, CL; freight,insurance, and taxes, CFIT; construction overhead, CO; and contractor engineering, CE.These factors are summed to give the total bare module factor. For the second option, theEntering Purchase Cost and Bare Module Factor dialog box appears, in which theequipment name, purchase cost, and bare module factor are entered. Bare module factorsfor a number of types of equipment are given in Table 16.11. Alternatively, for the thirdoption, the Entering Bare Module Cost dialog box appears. Here, just the equipmentname and bare module cost are entered.

Aspen IPE Specifications. When Aspen IPE is used to estimate purchase andinstallation costs for the entire plant, or a portion of the plant, click the Options entry inthe Menu and check the Allow IPE Entries box at the bottom of the dialog box thatappears. This produces the IPE Specifications subsection in the Equipment Costs section.Then left-click within the IPE Specifications subsection to produce the IPE Specifications

CD-17.8-5dialog box. The entries, Total Direct Materials and Labor Costs, Material and LaborG&A Overhead and Contractor Fees, Contractor Engineering Costs, and Indirect Costs,are obtained from the Aspen IPE Capital Estimate Report, as discussed in Section 16.7.Other costs can be entered (e.g., for pipe racks and sewers/sumps) under MiscellaneousInstallation Costs, if desired. The entries are summed and added to the Total BareModule Cost.

Total Permanent Investment

When left clicking within the Total Permanent Investment section, the DirectPermanent Investment dialog box appears. For each of the pertinent entries in Table16.9, the default entry can be altered. Either a percentage value or an absolute dollaramount is entered.

Working Capital

When left-clicking within the Working Capital section, the Working Capitaldialog box appears, on which the numbers of days are provided for the product, accountsreceivable, cash reserves, and accounts payable, as discussed in Section 17.3, withdefaults for each. If desired, additional entries can be made for any or all of the rawmaterials.

Utilities

When left-clicking within the Utilities section, the Utilities dialog box appears.An entry is provided for six default utilities in the spreadsheet: high pressure steam, lowpressure steam, process water, cooling water, natural gas, and electricity. Additionalutilities can be entered by pressing the Options button in the Menu, selecting the Utilitiestab, pressing the Add button to produce the Add a Utility dialog box, entering the name ofthe utility (e.g., medium pressure steam), and pressing the Add button. The additionalutilities appear in the Utilities dialog box. Then check the box for each utility in the

CD-17.8-6process and press the OK button, to produce a box into which its unit of measure isentered (e.g., kWhr for electricity). This produces a named utility dialog box in whichthe units of utility per unit of product is entered (e.g., lb high pressure steam/lb MCB), aswell as the price of the utility (e.g., $/kWhr). By pressing the next button, a similardialog box is produced for the next utility, until this information is entered for all utilitiesin the process. Representative prices for many utilities are listed in Table 17.1.

Byproducts

When left-clicking within the Byproducts section, the Byproducts dialog boxappears. As in the specification of raw materials, byproducts are added individually, witha specification of the unit of measure, the unit of byproduct per unit of product, and theprice per unit of byproduct.

Other Variable Costs

When left-clicking within the Other Variable Costs section, the General Expensesdialog box appears, which permits the specification of percentages of product salescharged for selling/transfer expenses, direct research, allocated research, administrativeexpenses, and management incentives compensation. The defaults shown are those in thecost sheet of Table 17.1 and discussed in Section 17.2.

Fixed Costs

The entries under Fixed Costs appear in six subsections: Operations,

Maintenance, Operating Overhead, Property Taxes and Insurance, Straight-lineDepreciation, and Depletion Allowance. When left-clicking within each subsection, theappropriate dialog box appears, in which the default entries can be replaced whendesired. Note that the default entries are those in Table 17.1. However, underOperations, entries must be made for (1) the number of operators per shift, (2) technical

CD-17.8-7assistance to manufacturing, and (3) control laboratory, for which see page 576 and Table17.1. If a depletion allowance applies, see pages 606-608 for estimating it.

Financial Information

In addition to the above entries, it is necessary to specify financial information for

calculation of the return on investment (ROI), the net present value (NPV) and theinvestor’s rate of return (IRR), also known as the discounted cash flow rate of return(DCFRR). To accomplish this, select the Options button in the Menu, and the FinancialInformation tab. Then enter the income tax rate, the cost of capital (for the NPVcalculation) and the inflation rate. Note that a general inflation rate can be specified,applicable to all operating costs, or by checking the Different Inflation Rate box, separateinflation rates can be specified for fixed and variable costs.

Running the Analysis and Creating a Report

To initiate the profitability analysis, on the Menu, press the Create Report buttonto produce the Create Report dialog box. After entering the Report Name, to which theword “Report” is automatically appended, and the directory into which the report file isto be stored (i.e., the report path), press the Create Report button. The results, which areplaced in an EXCEL report file, include sections on the Investment Summary, whichpresents cost estimates for all entries associated with the total permanent investment, theworking capital, and the total capital investment (i.e,, the entries shown in Table 16.9).Also included are sections on the Variable Costs at design capacity (not productionduring a specific year) of operations and for the Fixed Costs. These correspond to theentries shown in Table 17.1. Then, a section on the cash flows, and the elements thatcontribute to them, is displayed for each year during the life of the project. Finally, asection on the NPV and the IRR is provided. Each section is accessed by clicking on theappropriate tab at the bottom of the frame.

CD-17.8-8 It is also possible to have the ROI (during the third operating year) estimated andto carry out sensitivity studies. This is accomplished by pressing the Choose CustomAnalyses button, which produces the Custom Analyses dialog box, and checking the ROI(Third Year) entry. In addition, the IRR can be computed as a function of a singlevariable or as a function of two variables. These variables are the product price, variablecost, fixed cost, initial investment, and the rate of inflation.

Saving or Loading an Analysis

At any point when entering specifications or after completing an analysis, the

contents of the worksheet can be saved in a file. Alternatively, an existing file can beloaded into the spreadsheet. To accomplish this, on the spreadsheet Menu, press theSave/New button to produce the Save/New dialog box. To save a file, check the SaveCurrent Analysis button and press OK. On the Save As dialog box, enter a file name anda file path. To load a file, check the Load Existing Analysis button and press OK, whichproduces the Browse for Folder dialog box, within which the file is located. Note that theSave/New dialog box also permits the user to start a new analysis.

Having described the details of data entry into the spreadsheet, ProfitabilityAnalysis-1.0.xls, Example 17.32 is provided to illustrate its use for the MCB separationprocess.

Example 17.32 It is desired to carry out a profitability analysis for the monochlorobenzene (MCB) separation process using (a) purchase costs and bare module factors, (b) purchase and installation costs estimated by Aspen IPE. In Section 16.7, the latter estimates were computed, beginning with the ASPEN PLUS simulation in the file, MCB.bkp. Plant location is the Gulf Coast. The design time is estimated to be one year, the construction time at one year, and the total operating life of the project at 15 years. Assume that 5% of the total permanent investment is

CD-17.8-9allocated to engineering during the design year. The cost of capital is taken to be15% annually.

Fixed Costs Operations Operators per Shift: 1 (Assuming 5 Shifts) Direct Wages and Benefits: $30.00 per Operator Hour Direct Salaries and Benefits: 15.00% of Direct Wages and Benefits Operating Supplies and Services: 6.00% of Direct Wages and Benefits Technical Assistance to Manufacturing: $0.00 per year, for each Operator per Shift Control Laboratory: $0.00 per year, for each Operator per Shift

Maintenance Wages and Benefits: 3.50% of Total Depreciable Capital Salaries and Benefits: 25.00% of Maintenance Wages and Benefits Materials and Services: 100.00% of Maintenance Wages and Benefits Maintenance Overhead: 5.00% of Maintenance Wages and Benefits

Operating Overhead General Plant Overhead: 7.10% of Maintenance and Operations Wages and Benefits Mechanical Department Services: 2.40% of Maintenance and Operations Wages and Benefits Employee Relations Department: 5.90% of Maintenance and Operations Wages and Benefits Business Services: 7.40% of Maintenance and Operations Wages and Benefits

Property Taxes and Insurance

Property Taxes and Insurance: 2.00% of Total Depreciable Capital

Straight Line Depreciation

Direct Plant: 8.00% of Total Depreciable Capital, less1.18 times the Allocated Costs for Utility Plants and Related Facilities Allocated Plant: 6.00% of 1.18 times the Allocated Costs for Utility Plants and Related Facilities

IRR Analysis - Two Variable

May, 2003 Monochlorobenzene Separation Process

Product Prices vs Variable Costs

Variable Costs $17,689,500 $18,209,800 $18,730,100 $19,250,400 $19,770,600 $20,290,900 $20,811,200 $21,331,500 $21,851,800 $22,372,000 $22,892,300 $23,412,600 $ 0.46 24.98% 18.44% 11.26% 3.06% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.47 31.49% 25.51% 19.12% 12.14% 4.25% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.49 37.46% 31.87% 26.01% 19.77% 12.99% 5.38% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.50 43.01% 37.73% 32.24% 26.50% 20.40% 13.79% 6.43% Out of Range Out of Range Out of Range Out of Range Out of Range Product Prices

Fixed Costs Operations Operators per Shift: 1 (Assuming 5 Shifts) Direct Wages and Benefits: $30.00 per Operator Hour Direct Salaries and Benefits: 15.00% of Direct Wages and Benefits Operating Supplies and Services: 6.00% of Direct Wages and Benefits Technical Assistance to Manufacturing: $0.00 per year, for each Operator per Shift Control Laboratory: $0.00 per year, for each Operator per Shift

Maintenance Wages and Benefits: 3.50% of Total Depreciable Capital Salaries and Benefits: 25.00% of Maintenance Wages and Benefits Materials and Services: 100.00% of Maintenance Wages and Benefits Maintenance Overhead: 5.00% of Maintenance Wages and Benefits

Operating Overhead General Plant Overhead: 7.10% of Maintenance and Operations Wages and Benefits Mechanical Department Services: 2.40% of Maintenance and Operations Wages and Benefits Employee Relations Department: 5.90% of Maintenance and Operations Wages and Benefits Business Services: 7.40% of Maintenance and Operations Wages and Benefits

Property Taxes and Insurance

Property Taxes and Insurance: 2.00% of Total Depreciable Capital

Straight Line Depreciation

Direct Plant: 8.00% of Total Depreciable Capital, less1.18 times the Allocated Costs for Utility Plants and Related Facilities Allocated Plant: 6.00% of 1.18 times the Allocated Costs for Utility Plants and Related Facilities

IRR Analysis - Two Variable

May, 2003 Monochlorobenzene Separation Process

Product Prices vs Variable Costs

Variable Costs $17,689,500 $18,209,800 $18,730,100 $19,250,400 $19,770,600 $20,290,900 $20,811,200 $21,331,500 $21,851,800 $22,372,000 $22,892,300 $23,412,600 $ 0.46 17.98% 12.43% 6.18% -1.20% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.47 23.51% 18.52% 13.10% 7.04% -0.04% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.49 28.55% 23.94% 19.04% 13.74% 7.86% 1.06% Out of Range Out of Range Out of Range Out of Range Out of Range Out of Range $ 0.50 33.21% 28.88% 24.35% 19.54% 14.36% 8.64% 2.09% Out of Range Out of Range Out of Range Out of Range Out of Range Product Prices